Apparatuses and methods for hydrogen production

ABSTRACT

The present disclosure provides systems and methods for hydrogen production as well as apparatuses useful in such systems and methods. Hydrogen is produced by steam reforming of a hydrocarbon in a gas heated reformer that is heated using one or more streams comprising combustion products of a fuel in an oxidant, preferably in the presence of a carbon dioxide circulating stream.

CROSS-REFERENCE TO RELATED APPLICATIONS

The present application claims priority to U.S. Provisional PatentApplication No. 63/280,761, filed Nov. 18, 2021, U.S. Provisional PatentApplication No. 63/280,774, U.S. Provisional Patent Application No.63/280,786, filed Nov. 18, 2021, U.S. Provisional Patent Application No.63/280,793, filed Nov. 18, 2021, and U.S. Provisional Patent ApplicationNo. 63/423,301, filed Nov. 7, 2022, the disclosures of which areincorporated herein by reference in their entireties.

FIELD OF THE DISCLOSURE

The present disclosure provides for production of hydrogen. Moreparticularly, the disclosure provides apparatuses that are configuredfor use in hydrogen production and methods for producing hydrogen thatcan incorporate the apparatuses. The apparatuses and methods may beutilized for hydrocarbon reforming to produce hydrogen. The apparatusesand methods likewise can incorporate for oxy-fuel combustion in theproduction of hydrogen.

BACKGROUND

Hydrogen combustion for energy production emits only water and avoidsthe production of carbon dioxide, which takes place when hydrocarbonfuels are burned for energy production. Hydrogen has been described asthe fuel of the future as its widespread use replacing hydrocarbon fuels(and complemented by renewable energy production from wind and solarsystems) is the best route to achieving low to zero anthropogenicemission of carbon dioxide and thereby helping to address climatechange. Hydrogen can be used as a replacement for hydrocarbon fuels. Inparticular, hydrogen may replace natural gas in the pipelinedistribution network supplying fuel for domestic, commercial, andindustrial heating.

Hydrogen has been described as the perfect fuel when used with fuelcells for all types of vehicles, both road and rail. Battery poweredcars must be fueled from carbon free sources of power (i.e.,electricity) available on demand in very large amounts in order forelectric vehicles to be effective in reducing global carbon dioxideemissions. To this end, hydrogen can replace natural gas in bothexisting and new gas turbine combined cycle power generation plants.Hydrogen diluted with nitrogen and/or steam has been demonstrated as afuel on existing gas turbines by major manufacturers. It is thusexpected that implementation of hydrogen for power production will be akey to rapidly implementing the elimination of carbon dioxide fromelectric power generation without massive additional capital expenditureon new electric generation systems or costly carbon dioxide captureretrofits on existing systems, which result in significant degradationin generation efficiency. The key to the implementation of the hydrogenenergy economy is to devise a generic process that has very high thermalefficiency, which is defined as the ratio of the lower heating value(LHV) of the hydrogen product divided by the LHV of the hydrocarbon feedused for hydrogen generation. This must be coupled with a low capitalcost simultaneous, near 100% capture of the carbon dioxide produced inthe hydrogen generation process for sequestration. The process must alsobe suitable for the construction of very large production capacity unitswith low technical risk.

Hydrogen can be produced by electrolysis of water in pressurizedelectrolysis plants, but power consumption is very high, and theelectricity source must be carbon free to secure an environmentaladvantage. The additional oxygen that is formed as a by-product ofelectrolysis must also be usefully employed. The likely use forelectrolysis will be confined to dealing with excess power productionfrom renewable sources. Hydrogen production from natural gas and lighthydrocarbon liquid fuels utilizes reactions between steam or oxygen or acombination of these two processes. For simplicity of explanation,methane (CH₄) will be used as the specimen hydrocarbon in this document,although other hydrocarbons may be interchanged as appropriate. The mostwidely used process is steam natural gas catalytic reforming. Thereactions taking place are shown below in Equations 1 and 2.

CH₄+H₂O=CO+3H₂  Eq. 1

CO+H₂O=CO₂+H₂ (i.e., “shift” reactions)  Eq. 2

The catalytic steam methane reforming (SMR) reaction is highlyendothermic with a heat of reaction of 49.201 Kcal/gm mol. The heat issupplied by burning methane plus waste fuel gas in a radiant furnaceoperating at near atmospheric pressure, which heats an array ofthick-walled tubes filled with catalyst and which operate at pressuresof around 30 bar to 35 bar. Maximum allowable operating pressure isabout 35 bar. Typically, the reaction is carried out in an excess ofsteam, with the molar ratio of steam to natural gas being in the rangeof about 3 to about 4. The temperature of the mixture of methane andsteam enters the catalytic reactor at a temperature of about 400° C. toabout 600° C., and the reactor discharge temperature is generally in therange of about 800° C. to about 900° C. The reaction product of carbonmonoxide and hydrogen (i.e., “synthesis gas” or “syngas”) is then cooledand passed through one or more shift reactors where the carbon monoxidereacts with water in the presence of a catalyst per Equation 2 togenerate more hydrogen and shift the carbon monoxide to carbon dioxide.

To provide the necessary heat for the endothermic reaction, thick walledtubes are generally arranged in spaced apart rows and are heated byradiant heat from flames that are directed downwardly between the spacedapart rows of thick walled tubes from burners mounted in the roof of thecatalytic reactor. These flames are generated by the combustion of airand methane (or other hydrocarbon), and the resulting carbon dioxide andwater vapor, together with nitrogen from the air, are vented toatmosphere. The combustion gases in the vicinity of the top (inlet) ofthe catalytic reactor are typically at temperatures of about 1800° C. toabout 2500° C., and the combustion gases leave the catalytic reactor attemperatures of about 1000° C. to about 1100° C.

The pressure outside of the thick walled, catalyst filled tubes istypically only slightly below ambient, which explains why the maximumpressure within the tubes does not normally exceed about 35 bar at atypical maximum tube surface temperature of about 1050° C. A maximumpressure ratio of approximately 35:1 is typically considered prudent forreliable operation and long tube life. The thick walled tubes can becentrifugally cast from a high nickel alloy, such as HK40.

An alternative process for production of syngas from hydrocarbon feed isthe partial oxidation (POX) of natural gas using pure oxygen accordingto the reaction shown below in Equation 3, which can be followed by theshift reactions of Equation 2.

CH₄+0.5O₂=CO+2H₂  Eq. 3

The partial oxidation reaction is slightly exothermic at 8.527 Kcal/gmmol, but the reactor must operate at a discharge temperature of about1300° C. to about 1400° C. for the maximum conversion of the hydrocarbonfeed to occur with a reasonable residence time in the reactor. Operatingpressure is governed by reactor design and can be as high as about 100bar.

An auto-thermal reformer (ATR) may also be utilized for production ofsyngas, and this can comprise a POX burner operating with excess methaneplus added steam with the hot exhaust gas passing through a bed ofsteam/methane reforming catalyst where further hydrogen generation takesplace according to the reaction of Equation 1, and the product syngas(hydrogen plus carbon monoxide) is produced at a temperature of about1050° C. The high temperature syngas product from these processes iscooled in a steam generator, which produces the steam required for thereactions, but they all generate a very large amount of excess heat,which must be exported in the form of excess steam production orgenerated electric power.

In order to utilize this excess heat available from the very hightemperature of the hydrogen production processes, a two stage processhas been devised in which a first stage POX or ATR is operated in seriesor in a parallel configuration with a steam/hydrocarbon catalyticreformer—i.e., a gas-heated reformer (GHR). This is heated by the hightemperature syngas product from both stages so that the product outletsyngas stream entering the steam generator is reduced in temperature toabout 600° C., and the quantity of steam produced is only sufficient forthe process requirements for syngas generation. About 25% to 30% extrahydrogen can be generated from a fixed quantity of methane compared to asingle stage POX or ATR.

The production of pure hydrogen from the hot syngas leaving the steamgenerator involves conversion of carbon monoxide to hydrogen in one ormore shift reactors coupled with syngas cooling with the heat evolvedbeing used for boiler feed water and methane preheating. The crudehydrogen stream is processed in a multi-bed pressure swing adsorber(PSA), which produces a substantially pure, pressurized hydrogen productand a low pressure waste gas stream. The carbon dioxide present in thecrude hydrogen stream can be removed from either the PSA feed or the PSAwaste gas streams by processing in a variety of different ways includingcarbon dioxide removal using chemical solvents, such as MDEA, orphysical solvents, such as Selexol™. A system described in U.S. Pat. No.8,900,355 separates carbon dioxide by condensation at a temperatureclose to the carbon dioxide solidification point where the partialpressure of carbon dioxide is minimized. The uncondensed gas can then berecycled to the syn-gas generation system. Each of the methods describedcan result in the removal of at least 90% and preferably near 100% ofthe carbon dioxide derived from carbon in the methane feed gas. The mostefficient processes currently available for hydrogen generation withnear 100% carbon dioxide capture are the POX plus GHR and the ATR plusGHR where all of the carbon dioxide derived from the total methane feedis present in the pressurized crude hydrogen stream following the shiftreactors and coolers.

The steam/methane catalytic hydrogen system (SMR) has the advantage ofcatalytically oxidizing the methane with water to form the hydrogenproduct and the carbon dioxide by-product so that no added oxygen isrequired. The disadvantage of the current SMR system is that carbondioxide must be removed by a combination of removal from the shiftedsyngas using chemical and/or physical methods, and the PSA gas can thenbe used as fuel. Alternatively, a large quantity of methane and all ofthe PSA waste gas containing the entire carbon dioxide product may beused as fuel gas in the reformer furnace to provide the very large heatof reaction plus the preheat for the reaction products, which means thatcarbon dioxide must be removed from the near atmospheric pressure stackgas at a concentration of about 12%. Removing large volumes of carbondioxide at near atmospheric pressure is very costly and also reducesoverall process efficiency. Since the furnace acts as a radiant heattransfer system, the exit combustion gases are at temperatures typicallyin the range of about 1000° C. to about 1100° C., which requires a largeheat recovery heat exchange unit that preheats methane plus water feedto the catalyst and generates a large quantity of medium pressureby-product steam.

SUMMARY OF THE DISCLOSURE

The present disclosure relates to hydrogen production methods,individual pieces of equipment or apparatuses that are useful forhydrogen production, and combinations of pieces of the equipment orapparatuses that together can define systems, units, or plantsconfigured for hydrogen production. The hydrogen production of thepresent disclosure can be carried out so that produced hydrogen can beisolated with increased purity of product and increased processefficiency through appropriate combinations of system components andsystem operational procedures. The hydrogen production can furtherexhibit increased efficiency through use of oxy-fuel heating that notonly provides process heat but also provides an integrated feature forprocessing waste gases back through the process to utilize potentialheat of combustion remaining in the waste gas. Oxy fuel heating producessubstantially only carbon dioxide and steam combustion products, whichallows for the separation of pressurized carbon dioxide forsequestration following separation condensed water without the need fora separate carbon dioxide removal system. The hydrogen production canalso improve process efficiency by allowing for relatively highoperational pressures such that by-product carbon dioxide can becaptured at increased pressures that simplify removal thereof andplacement for delivery for sequestration or other use of the carbondioxide. The hydrogen production can additionally exhibit improvedprocess efficiency and reduction materials costs through utilization ofspecifically chosen component parts of the unit/system/plant. This caninclude, for example, new and useful gas heated reforming (GHR) reactordesigns and/or ion transport membrane (ITM) technologies for directproduction of heat of combustion with oxygen separation from air.

The disclosure thus provides clean hydrogen technology that increaseshydrogen generation efficiency, reduces capital expenditure (CAPEX),reduces design complexity, and maintains near 100% carbon capture.

In one or more embodiments, the present disclosure can provide oxy-fuelheated, hydrogen production systems. These systems can be formed from avariety of combinations of components as described herein. In someembodiments, an example oxy-fuel heated, hydrogen production systems cancomprise:

a reforming reactor arranged to receive a stream comprising ahydrocarbon and water through a first inlet and separately receive astream of a heating fluid through a second inlet, the reactor includinga catalyst component effective for catalyzing a reaction between thehydrocarbon and the water to form a synthesis gas stream comprising atleast hydrogen and carbon monoxide, and the reactor including asynthesis gas outlet arranged for exit of the synthesis gas stream fromthe reforming reactor;

an oxy-fuel combustor arranged to receive a fuel, an oxidant, and astream comprising predominately carbon dioxide and comprising acombustor outlet for exit of a combustion product stream from theoxy-fuel combustor, the oxy-fuel combustor being configured to combustat least a portion of the fuel with oxygen from the oxidant to formcarbon dioxide and water, which is combined with the stream comprisingpredominately carbon dioxide to form the combustion product stream;

a hydrogen isolation unit arranged to receive at least a portion of thesynthesis gas stream, and provide at least part of the hydrogen from thesynthesis gas stream as a substantially pure hydrogen product stream;and

a purification unit arranged to receive at least a portion of thecombustion product stream and output a stream of substantially purecarbon dioxide, the purification unit also being arranged to deliver atleast a portion of the substantially pure carbon dioxide as the streamcomprising predominately carbon dioxide;

wherein the reforming reactor and the oxy-fuel combustor arefunctionally configured so that at least part of the combustion productstream is provided through the second inlet of the reforming reactor asthe stream of the heating fluid.

In further embodiments, an oxy-fuel heated, hydrogen production systemcan be further defined in relation to any one or more of the followingstatements. These following statements are intended to be combinable inany number and order, and it is understood that the express listing ofthese statements provides indication that each of the possiblecombinations are identifiable in light of the following statements asread in light of the full disclosure provided herein.

The reforming reactor can comprise a pressure containment vessel and atleast one set of concentrically arranged tubes positioned within thepressure containment vessel, each of the at least one set ofconcentrically arranged tubes comprising: an outer catalyst tube; aninner reaction product gas tube; and catalyst material positioned withina space defined between an inside surface of the outer catalyst tube andan outside surface of the inner reaction product gas tube.

The at least one set of concentrically arranged tubes positioned withinthe pressure containment vessel can be arranged vertically so that anupper end of the at least one set of concentrically arranged tubesdefines a hot end where the reforming reactor operates with a highertemperature, and a lower end of the at least one set of concentricallyarranged tubes defines a cold end where the reforming reactor operateswith a lower temperature, relative to the hot end.

The reforming reactor further can comprise an upper tube sheet that isarranged to functionally align with the outer catalyst tube, and a lowertube sheet that is arranged to functionally align with the innerreaction product gas tube.

The reforming reactor can be arranged so the first inlet opens into aspace defined between the upper tube sheet and the lower tube sheet.

The reforming reactor can be arranged so that the stream comprising ahydrocarbon and water entering through the first inlet passes upwardly,from the cold end toward the hot end, through the space within which thecatalyst material is positioned.

The reforming reactor can be arranged so that the synthesis gas outletis positioned at a level of the reforming reactor that is below aposition of the first inlet.

The reforming reactor can be arranged so that the synthesis gas outletis positioned below the lower tube sheet.

A bottom of the lower tube sheet and a bottom portion of the pressurecontainment vessel can define a collection space for the synthesis gasstream, which proceeds downwardly from the hot end through an inner boreof the inner reaction product gas tube.

The second inlet of the reforming reactor can be configured to receivethe heating fluid in an arrangement so that the heating fluid providesheat to the outer catalyst tube.

The arrangement can be such that the heating fluid entering the secondinlet of the reforming reactor contacts the hot end of the at least oneset of concentrically arranged tubes and flows downwardly around anouter surface of the outer catalyst tube toward a second outlet throughwhich the heating fluid leaves the reforming reactor.

The second outlet can be positioned at a level of the reforming reactorthat is above a position of the first inlet.

The reforming reactor further can comprise a surrounding tube positionedaround the at least one set of concentrically arranged tubes, thesurrounding tube being arranged to form heating space relative to the atleast one set of concentrically arranged tubes and define a flow path ofthe heating fluid through the heating space.

The reforming reactor further can comprise a plurality of bafflesattached to an inner surface of the pressure containment vessel andarranged to direct flow of the heating fluid for contact with the atleast one set of concentrically arranged tubes.

The upper end of the at least one set of concentrically arranged tubescan define a filling tube with a removable plug.

The removable plug can be configured to provide biased force toward thecatalyst within the at least one set of concentrically arranged tubes.

An outer surface of the at least one set of concentrically arrangedtubes can comprise a plurality of fins configured to facilitate heattransfer between the heating fluid and the at least one set ofconcentrically arranged tubes.

At least a portion of internal surfaces of the reforming reactor thatare exposed to a partial pressure of carbon monoxide at operatingtemperatures where a Bouduard reaction occurs can be protected frommetal dusting corrosion by the presence of a protective coating or alayer of internal insulation.

The space defined between the inside surface of the outer catalyst tubeand the outside surface of the inner reaction product gas tube that isfilled with catalyst can define a section having a length about 6 metersto about 18 meters.

The oxy-fuel heated, hydrogen production system further can comprise atleast one shift reactor configured to convert at least a portion of thecarbon monoxide in the synthesis gas from the reforming reactor tocarbon dioxide and output a shift stream comprising at least hydrogen,carbon dioxide, and waste gas.

The hydrogen isolation unit can comprise an inlet arranged to receivethe shift stream, output a pressurized stream of substantially purehydrogen, and output a stream comprising at least part of the waste gas.

The hydrogen isolation unit can comprise a hydrogen multi-bed pressureswing adsorber (PSA) configured to output the pressurized stream ofsubstantially pure hydrogen and output the stream comprising at leastpart of the waste gas.

The PSA can be configured with a hydrogen recycle line arranged to sendpart of the pressurized stream of substantially pure hydrogen back tothe inlet of the PSA.

The hydrogen isolation unit further can comprise at least one compressorarranged to receive and compress at least a portion of the streamcomprising at least part of the waste gas and output a compressed wastegas stream.

The hydrogen isolation unit further can comprise a membrane gasseparator having an inlet arranged to receive the compressed waste gasstream, and wherein the membrane gas separator is configured to separatethe compressed waste gas stream into a pressurized retentate waste gasstream and a hydrogen-enriched permeate stream.

The membrane gas separator can comprise an inlet arranged to receive astream of substantially pure carbon dioxide for passage through apermeate side of a membrane in the membrane gas separatorcounter-current to the compressed waste gas stream.

The hydrogen isolation unit further can comprise a recirculation linethrough which the hydrogen-enriched permeate stream is passed back tothe inlet of the PSA.

The oxy-fuel heated, hydrogen production system further can comprise aline through which at least part of the pressurized retentate waste gasstream is passed to the oxy-fuel combustor.

The oxy-fuel heated, hydrogen production system further can comprise agas turbine.

The oxy-fuel heated, hydrogen production system further can comprise aline through which at least a portion of the pressurized stream ofsubstantially pure hydrogen is passed to the gas turbine.

The oxy-fuel heated, hydrogen production system further can comprise anammonia synthesis unit.

The oxy-fuel heated, hydrogen production system further can comprise aline through which at least a portion of the pressurized stream ofsubstantially pure hydrogen is passed to the ammonia synthesis unit.

The oxy-fuel heated, hydrogen production system further can comprise apower producing turbine arranged to receive at least a portion of thesynthesis gas stream and expand said stream for power production.

The oxy-fuel combustor can comprise an outer combustor shell and acombustor liner that defines internally a combustion chamber.

The oxy-fuel combustor can be arranged to receive at least part of thestream comprising predominately carbon dioxide through the combustorliner.

The oxy-fuel combustor can be arranged to receive a first part of thestream comprising predominately carbon dioxide into a reaction zone ofthe combustion chamber and to receive a second part of the streamcomprising predominately carbon dioxide into a dilution zone of thecombustion chamber.

The oxy-fuel combustor can comprise an ion transport membrane (ITM)combustor.

The ITM combustor can comprise an oxygen ion transport diffusionmembrane separating an air side of the ITM combustor from a fuel side ofthe ITM combustor.

The oxygen ion transport diffusion membrane can be effective to drawoxygen from air passing through the air side of the ITM combustor intothe fuel side of the ITM combustor for combustion of fuel passed throughthe fuel side of the ITM combustor.

The oxy-fuel heated, hydrogen production system can comprise a pluralityof ITM combustors.

The oxy-fuel heated, hydrogen production system further can comprise aheat exchanger arranged to receive at least a portion of the heatingfluid after the heating fluid exits the reforming reactor and configuredto transfer heat from the heating fluid to one or more further streams.

The one or more further streams to which the heat is transferred fromthe heating fluid can include one or more of the fuel that is receivedby the oxy-fuel combustor, the oxidant that is received by the oxy-fuelcombustor, the stream comprising predominately carbon dioxide that isreceived by the oxy-fuel combustor, and the stream comprising thehydrocarbon and water that is received by the reforming reactor.

The oxy-fuel heated, hydrogen production system further can comprise apurification unit arranged to receive the heat fluid after leaving theheat exchanger and configured to output the stream comprisingpredominately carbon dioxide.

The oxy-fuel heated, hydrogen production system further can comprise acompressor arranged to receive the stream comprising predominatelycarbon dioxide leaving the purification unit and configured to compressthe stream comprising predominately carbon dioxide to a pressuresuitable for input to the oxy-fuel combustor.

These and other features, aspects, and advantages of the disclosure willbe apparent from a reading of the following detailed descriptiontogether with the accompanying drawings, which are briefly describedbelow. The disclosure includes any combination of elements, components,and features that are described herein, regardless of whether suchelements, components, and features are expressly combined in a specificembodiment description herein. This disclosure is intended to be readholistically such that any separable features, components, or elementsof the disclosure, in any of its various aspects and embodiments, shouldbe viewed as intended to be combinable unless the context clearlydictates otherwise.

BRIEF DESCRIPTION OF THE FIGURES

Having thus described the disclosure in the foregoing general terms,reference will now be made to the accompanying drawings, which is notnecessarily drawn to scale, and which should be viewed as illustratingexample embodiments of the presently disclosed subject matter.

FIG. 1 is a flowchart illustrating steps in an oxy-fuel hydrogenproduction process according to example embodiments of the presentdisclosure.

FIG. 2A is a partial cross-sectional illustration of an example designfor a gas heated reforming reactor according to embodiments of thepresent disclosure.

FIG. 2B is a transecting view of the partial cross-sectionalillustration of FIG. 2A showing the arrangement of the tubes in the gasheated reforming (GHR) reactor.

FIG. 3A is a partial cross-sectional illustration of another exampledesign for a gas heated reforming reactor (GHR) according to embodimentsof the present disclosure.

FIG. 3B is a transecting view of the partial cross-sectionalillustration of FIG. 3A showing the arrangement of the tubes in the gasheated reforming reactor.

FIG. 4 is a partial cross-sectional illustration providing a moredetailed arrangement of a single reactor tube within a gas heatedreforming reactor according to embodiments of the present disclosure.

FIG. 5 is a flow diagram illustrating at least a portion of thecomponents useful in an oxy-fuel hydrogen production process accordingto example embodiments of the present disclosure.

FIG. 6 is a flow diagram illustrating at least a portion of thecomponents useful in a hydrogen isolation unit, which can beparticularly useful in oxy-fuel hydrogen production processes accordingto example embodiments of the present disclosure.

FIG. 7 is flow diagram illustrating an oxy-fuel hydrogen productionsystem according to example embodiments of the present disclosure usefulfor carrying out oxy-fuel hydrogen production processes according toexample embodiments of the present disclosure.

FIG. 8 is a flow diagram illustrating an oxy-fuel hydrogen productionsystem according to example embodiments of the present disclosure usefulfor carrying out oxy-fuel hydrogen production processes according toexample embodiments of the present disclosure.

FIG. 9 is a flow diagram illustrating an oxy-fuel hydrogen productionsystem according to example embodiments of the present disclosure usefulfor carrying out oxy-fuel hydrogen production processes according toexample embodiments of the present disclosure including the use of a gasturbine exhaust as an oxidant to combust waste gas as fuel to provideheat for a circulating stream of predominately carbon dioxide for use asa heating fluid in a reforming reactor.

FIG. 10 is a flow diagram illustrating an ion transport membrane (ITM)combustor arrangement useful in oxy-fuel hydrogen production systemsaccording to example embodiments of the present disclosure useful forcarrying out oxy-fuel hydrogen production processes according to exampleembodiments of the present disclosure.

DETAILED DESCRIPTION OF THE DISCLOSURE

The present subject matter will now be described more fully hereinafterwith reference to exemplary embodiments thereof. These exemplaryembodiments are described so that this disclosure will be thorough andcomplete, and will fully convey the scope of the subject matter to thoseskilled in the art. Indeed, the subject matter can be embodied in manydifferent forms and should not be construed as limited to theembodiments set forth herein; rather, these embodiments are provided sothat this disclosure will satisfy applicable legal requirements. As usedin the specification, and in the appended claims, the singular forms“a”, “an”, “the”, include plural referents unless the context clearlydictates otherwise.

The present disclosure provides for improved manners of hydrogenproduction and processes, systems, and equipment that can individually,or in combination, exhibit the improvements in the production ofhydrogen. Several embodiments are provided herein, and the severalembodiments are described individually only for ease of disclosure andease of understanding. The several embodiments, however, are expresslyintended to be useful individually or in any combination of the severalembodiments. It is understood that each embodiment provides improvementsin hydrogen production arising from the specific features of theindividual embodiment. The individual embodiments arise from recognitionof shortcomings in the existing methods and equipment used for hydrogenproduction, and each individual embodiment thus provides a usefulimprovement and advantage in hydrogen production. The improvements andadvantages can be multiplied through combinations of the individualembodiments, and the unique features of each embodiment are evidencethat the improvements achieved with the combinations of the embodimentsare not an expected, cumulative effect but are rather synergisticeffects arising from the various combinations of the individualembodiments.

In some embodiments, the present disclosure provides for oxy-fuelhydrogen production systems and methods that can use a carbon dioxideremoval system upstream of a hydrogen PSA apparatus or unit, such aswith a carbon dioxide condensation system, which can be expedient sincethere is no need to remove all of the carbon dioxide in the streamentering the PSA. In some embodiments, the present disclosure providesfor oxy-fuel hydrogen production systems and methods that include ahydrogen isolation unit comprising a hydrogen PSA plus a membranesystem, which enables use of a single PSA while still recoveringsignificant proportions of the hydrogen present in the gas stream beingtreated, even greater than 98% hydrogen recovery in certain embodiments.In some embodiments, the present disclosure provides for oxy-fuelhydrogen production systems and methods that can use an Ion TransportMembrane (ITM) oxygen supply system in place of a cryogenic AirSeparation Unit (ASU), and this can enable efficient combustion ofhydrocarbon fuel (and PSA waste gas, in certain embodiments) within theITM unit to produce a heating fluid stream for a GHR without therequirement for a separate combustor. In some embodiments, the presentdisclosure provides for oxy-fuel hydrogen production systems and methodsthat can incorporate a new gas heated reforming (GHR) reactorspecifically configured to make use of an oxy-fuel combustion productstream as a heating fluid.

A general flow diagram of a process for hydrogen production from ahydrocarbon source stream is illustrated in FIG. 1 . As seen therein,the hydrocarbon of choice is reacted at step 1000 with steam in thepresence of a catalyst to produce syngas. It is understood that use ofthe term “syngas” herein is referencing synthesis gas, which ispredominately carbon monoxide (CO) and hydrogen (H₂) but which also maycontain small amounts of additional components. Since a focus of thepresent disclosure to hydrogen production, the disclosure addressedmainly the hydrogen product that is originally present in the synthesisgas but also the carbon monoxide since it can be converted to additionalhydrogen and also carbon dioxide in the shift reactions also describedherein. Reference to hydrogen and carbon monoxide when discussing thesynthesis gas, or syngas, is thus not meant to exclude other componentsthat may be present in the syngas in typically small amounts, unless thecontext of the discussion clearly intends to only address the hydrogenand/or the carbon monoxide portions of the syngas.

Process heat can be provided at step 1005 and may be provided from avariety of sources. Syngas produced in step 1000 is processed at step1010 in one or more shift reactors to convert carbon monoxide to carbondioxide (CO₂) and produce additional hydrogen as well. The shiftedsyngas stream is then processed at step 1020 in a pressure swingabsorber in order to isolate hydrogen, which is taken for further usesat step 1025. A waste gas stream is likewise produced and can comprise,for example additional hydrogen, CO, carbon dioxide, and unreactedhydrocarbon. This waste gas is processed at step 1030 for additionaluses. This can include separation of at least a portion of the waste gasinto individual components for capture, and the individual componentsmay be removed as individual waste streams at step 1035. In someembodiments, at least a portion of the waste gas can be sent in step1040 to an oxy-fuel combustion component to be combusted withhydrocarbon fuel. The combustion products then can form at least a partof the process heating noted in step 1005 above.

Hydrogen production according to the present disclosure can proceed viathis general process, which process is modified by any single embodimentof the present disclosure or by any combination of two embodiments,three embodiments, or more embodiments of the preset disclosure. In anexample embodiment, a hydrogen production can be carried out using asingle stage pressurized catalytic steam plus methane syngas generationreactor including a supply of the required heat and can utilize asimple, low cost hydrogen separation system that is effective to achievehigh purity hydrogen isolation with up to 100% carbon dioxide capturefrom waste gas streams, which waste gas streams can be provided at ahigh pressure typically required for hydrogen production.

The present disclosure particularly can utilize a gas heated reformer(GHR) that is configured to operate under conditions not previouslyattainable and thereby provide for greatly improved efficiency ofhydrogen production, reduced capital costs for hydrogen production,reduced operational costs for hydrogen production, and simplifiedoperational procedures for hydrogen production. The GHR arrangementsdescribed herein by various example embodiments can be utilized invarious hydrogen production processes, and the modification of suchprocesses to utilize one or more of the present GHR arrangements can, byitself, provide the desired outcomes otherwise described herein forhydrogen production. It is understood, however, that the presentdisclosure also encompasses further embodiments wherein theimplementation of one or more of the GHR arrangements described hereincan be combined with one or more of the additional process improvementsand/or apparatuses described herein.

The disclosure particularly provides a heating unit that achieves veryhigh efficiency of methane conversion to hydrogen in the form of a gasheated reforming reactor (GHR) that provides for production of hydrogenat significantly higher pressures than previously used in the art due toits design features and its ability to control conversion of methane(and/or other hydrocarbon fuels) to syngas. The GHR can utilize a higherconversion pressure and can efficiently recycle fuel gas containingunconverted hydrocarbon back to an associated heating component,preferably an oxy-fuel burner.

The first optional arrangement of a GHR 100 is provided in FIG. 2A. TheGHR 100 comprises an outer vessel 105, which can be a pressure vessel.The outer vessel 105 extends from a first end 106, which is a lower orbottom end, to a second end 107, which is an upper or top end. Withinthe outer vessel 105 is one or more tube assemblies, and the outervessel preferably includes a plurality of the tube assemblies. A singletube assembly is described following, but it is understood that thedescription can apply to any one, two, or more of the tube assembliesthat are present within the outer vessel 105.

An individual tube assembly within the GHR 100 comprises an assemblyouter tube 120 and an assembly inner tube 130. The assembly outer tube120 may be referenced herein as an outer catalyst tube since, asdiscussed below, a catalyst material can be retained within the assemblyouter tube 120 along with the concentrically arranged assembly innertube 130. An assembly inner tube 130 may be referenced herein as aninner reaction product gas tube since the syngas formed by the catalyticreaction of hydrocarbon and steam while passing through the catalystwill be produced downwardly through the assembly inner tube 130. Asillustrated, the tubes are configured to have a cross-sectional shapethat is substantially round; however, other cross-sectional shapes arealso encompassed by the present disclosure, such as elliptical,rectangular, square, and the like. The assembly outer tube 120 andassembly inner tube 130 are concentrically arranged such that a space125 exists between the tubes (see FIG. 2B).

The assembly inner tube 130 is configured as a central syngas productoutlet tube. The vessel 105 is arranged to receive heating gas at theupper end 107, and the upper portion of the GHR can thus becharacterized as the hot end of the GHR 100 while the lower end 106 canbe characterized as the cold end of the GHR 100, and the tube assemblieswithin the GHR 100 can likewise be characterized as having hot ends andcold ends in relation to the location thereof relative to the hot end ofthe GHR 100 and the cold end of the GHR 100. The terms “hot” and “cold”are thus used to define relative temperature conditions between theopposing upper end 107 and lower end 106. The hot end of the assemblyouter tube 120 can comprise a removable cap 127 or hot end cover. Theterm “cap” is not intended to limit the structure of the element, andthe element need only be configured to removably engage the hot end ofthe assembly outer tube 120 to close off the hot end thereof. The cap127 may engage the outer surface of the assembly outer tube 120 or theinner surface of the assembly outer tube 120 to provide the closingfunction. To this end, the hot end of the assembly outer tube 120 can beconfigured as a filling head 126, and this can be a shape or otherarrangement that allows for ease of addition of catalyst to the space125 between the assembly outer tube 120 and the assembly inner tube 130.The cap 127 also can be functional to cause syngas that is formed asreaction materials pass upwardly between the assembly inner tube 130 andthe assembly outer tube 120 to turn and move downwardly through theinterior of the assembly inner tube 130. An upper portion of theassembly inner tube 130 can include a top plug 123 that is configured tosubstantially prevent passage of any catalyst material therethroughwhile allowing passage of gas. The top plug 123 can thus be made of aporous material, such as stainless steel or nickel alloy, with anaverage pore size of less than 1 mm, less than 0.5 mm, or less than 0.2mm, such as about 0.01 mm to about 0.95 mm, about 0.02 mm to about 0.75mm, or about 0.05 mm to about 0.5 mm.

The space 125 between the assembly outer tube 120 and the assembly innertube 130 is at least partially filled with a catalyst for reforming thesteam plus hydrocarbon that is processed into the GHR 100. The catalystpreferably is granular, particulate, or otherwise in a suitable shapeand size to provide the necessary surface area for catalyzing thereforming reactions. In some embodiments, the space 125 can be definedas an annular space. More particularly, the space 125 is defined by aninside surface 121 of the assembly outer tube 120 and an outside surface131 of the assembly inner tube 130. The space 125 between the assemblyouter tube 120 and the assembly inner tube 130 can define one or moresections that are filled with the catalyst. In some embodiments, thecatalyst filled section can extend up substantially to the filling head126 or to the cap 127. A bottom of the catalyst filled section can bedefined by a bottom plug 133 that is configured to substantially preventpassage of any catalyst material therethrough while allowing passage ofgas. The bottom plug 133 can thus be made of a porous material, such asstainless steel or nickel alloy, with an average pore size of less than1 mm, less than 0.5 mm, or less than 0.2 mm, such as about 0.01 mm toabout 0.95 mm, about 0.02 mm to about 0.75 mm, or about 0.05 mm to about0.5 mm. The length of the catalyst filled section can be about 5 metersto about 24 meters, about 7 meters to about 20 meters, or about 10meters to about 18 meters.

The assembly inner tube 130 can have an outer diameter in the range ofabout 15 mm to about 40 mm, about 18 mm to about 35 mm, or about 20 mmto about 30 mm. In an example embodiment, assembly inner tube 130 canhave an outer diameter of 25.4 mm. The assembly outer tube 120 can havean inner diameter of about 45 mm to about 120 mm, about 50 mm to about110 mm, or about 60 mm to about 100 mm. In an example embodiment, theassembly outer tube 121 can have an inner diameter of 76 mm. Theassembly outer tube 120 and the assembly inner tube 130 each canindependently be made from corrosion resistant materials, such asstainless steel or high nickel alloy, and can have a wall thickness of,for example, about 0.5 mm to about 5 mm, about 0.8 mm to about 4 mm, orabout 1 mm to about 3 mm.

The tube assemblies are supported by tube plates, which are arranged assubstantially flat sheets with holes positioned and sized toappropriately meet ends of the individual tubes in the tube assemblies.As illustrated in FIG. 2A, an upper tube sheet 129 is arranged tofunctionally align with the assembly outer tube 120, and a lower tubesheet 139 is arranged to functionally align with the assembly inner tube130. The upper tube sheet 129 may be characterized as a feed gas inlettube sheet, and the lower tube sheet 139 may be characterized as asyngas outlet tube sheet. Preferably, the functional alignment includeswelded connections between the tube sheets and the tube or otherarrangements to provide substantially sealed sections defined in part bythe tube sheets and in part by the outer vessel 105. For example, a feedgas inlet distribution space 108 can be defined between a bottom surfaceof the upper tube sheet 129 and a top surface of the lower tube sheet139. As a further example, the lower end 106 of the GHR can define anoutlet collection space between a bottom surface of the lower tube sheet139 and the inner surface of the bottom of the vessel 105.

The vessel 105 further includes an inlet 150 configured to receive areaction feed stream 151 comprising steam plus hydrocarbon. A syngasproduct stream 153 leaves the GHR 100 through an outlet 152. The vessel105 also includes an inlet 157 arranged to receive a heating fluidstream 156. A series of baffles 155 can be arranged on an interiorsurface of the wall defining the vessel 105, and these can be functionalto provide a baffled multi-pass shell side flow path for the heatingfluid through the vessel 105. The heating fluid flows downwardly in thevessel 105 and exits through an outlet 159 as return heat fluid stream158.

In order to minimize the design temperature of the vessel 105 definingthe GHR arrangement 100, and to allow low cost alloys to be used forconstruction, internal insulation can be used. The inner surface of theassembly inner tube 130, the inner surface of the assembly outer tube120, and the surfaces of the upper tube sheet 129 and lower tube sheet139 will be exposed to high partial pressures of carbon monoxide attemperatures where the Bouduard reaction will take place with thepotential to cause metal dusting corrosion. To address such reactions,corrosion resistant alloys, such as Specialty Metals alloy 693, may beused, and/or exposed components may be coated with a plasma sprayedimpervious oxide layer, such as alumina, and/or exposed surfaces may becovered with impervious internal insulation.

The assembly outer tube 120 can have finned outer surfaces to promoteheat transfer. In embodiments with heating fluid passing across thetubes, radial fins may be preferred while, in other embodiments,longitudinal fins may be preferred. The GHR 100 is preferentiallyarranged vertically (i.e., with a vertically aligned longitudinal axis)so that the hot ends are the ends including the filling heads 126 andcaps 127. This arrangement provides the ability to inject the heatingfluid 156 to flow downward through the vessel 105 to heat the tubeassemblies while also providing the filling heads 126 and the caps 127at the top of the GHR to facilitate the filling and emptying of catalystin the assembly outer tubes 120. Addition and/or removal of catalyst maybe carried out by removal of a removable top section 109 of the vessel105. This top section 109 may be lifted off for servicing of the GHR 100and replacement of catalyst in the tube assemblies, and the top section109 can be replaced for operation of the GHR 100.

A second optional arrangement of a GHR 200 according to an exampleembodiment of the present disclosure is illustrated in FIG. 3A. The GHR200 has the same overall layout as the first GHR arrangement but alsohas the addition of a third surrounding tube. With reference to FIG. 3A,the GHR 200 again comprises an outer vessel 205, which can be a pressurevessel. The outer vessel 205 extends from a first end 206, which is alower or bottom end, to a second end 207, which is an upper or top end.Within the outer vessel 205 is one or more tube assemblies, and theouter vessel preferably includes a plurality of the tube assemblies. Asingle tube assembly is described following, but it is understood thatthe description can apply to any one, two, or more of the tubeassemblies that are present within the outer vessel 205. Moreover, it isunderstood that description above of the materials, arrangements, andconfigurations of the first arrangement of a GHR 100 can equally applyto the second arrangement of a GHR 200 that is further discussed below.

An individual tube assembly within the GHR 200 comprises an assemblyouter tube 220 and an assembly inner tube 230. As illustrated, the tubesare configured to have a cross-sectional shape that is substantiallyround; however, other cross-sectional shapes are also encompassed by thepresent disclosure, such as elliptical, rectangular, square, and thelike. The assembly outer tube 220 and assembly inner tube 230 areconcentrically arranged such that a space 225 exists between the tubes(see FIG. 3B).

The assembly inner tube 230 is configured as a central syngas productoutlet tube. The vessel 105 is arranged to receive heating gas at thehot, upper end 207 of the GHR. The hot end of the assembly outer tube220 defines a filling head 226 and can comprise a removable cap 227 orhot end cover having the same configurations as already described above.The space 225 between the outer assembly tube 220 and the inner assemblytube 230 is again at least partially filled with a catalyst forreforming the steam plus hydrocarbon that is processed into the GHR 100.In GHR 200, the space 125 again is defined by an inside surface 221 ofthe assembly outer tube 220 and an outside surface 231 of the assemblyinner tube 230. The space 225 between the assembly outer tube 220 andthe assembly inner tube 230 can define one or more sections that arefilled with the catalyst, as already described above. An upper portionof the assembly inner tube 130 can include a top plug 223 that isconfigured to substantially prevent passage of any catalyst materialtherethrough while allowing passage of gas. A bottom of the catalystfilled section can be defined by a bottom plug 233 that is configured tosubstantially prevent passage of any catalyst material therethroughwhile allowing passage of gas. The top plug 223 and the bottom plug 233can thus be made of a porous material, such as stainless steel or nickelalloy, with an average pore size of less than 1 mm, less than 0.5 mm, orless than 0.2 mm, such as about 0.01 mm to about 0.95 mm, about 0.02 mmto about 0.75 mm, or about 0.05 mm to about 0.5 mm. The assembly innertube 230 and assemble outer tube 220 can again have dimensions asalready described. The tube assemblies are supported by tube plates, andan upper tube sheet 229 is arranged to functionally align with theassembly outer tube 220, and a lower tube sheet 239 is arranged tofunctionally align with the assembly inner tube 230. The upper tubesheet 229 may be characterized as a feed gas inlet tube sheet, and thelower tube sheet 239 may be characterized as a syngas outlet tube sheet.Preferably, the functional alignment includes welded connections betweenthe tube sheets and the tube or other arrangements to providesubstantially sealed sections defined in part by the tube sheets and inpart by the vessel. For example, a feed gas inlet distribution space 208can be defined between a bottom surface of the upper tube sheet 229 anda top surface of the lower tube sheet 239. As a further example, thelower end 206 of the GHR can define an outlet collection space between abottom surface of the lower tube sheet 239 and the inner surface of thebottom of the vessel 205.

The vessel 205 further includes an inlet 250 configured to receive areaction feed stream 251 comprising steam plus hydrocarbon. A syngasproduct stream 253 leaves the GHR 200 through an outlet 252. The vessel205 also includes an inlet 257 arranged to receive a heating fluidstream 256. The heating fluid flows downwardly in the vessel 205 andexits through an outlet 259 as return heat fluid stream 258. The vessel205 again also will include construction, insulation, and/or coatings asalready described above to minimize the design temperature preventcorrosion or other fouling of parts arising from, for example, theBouduard reaction. The assembly outer tube 220 can have finned outersurfaces to promote heat transfer. In such embodiments, longitudinalfins are preferred while, in other embodiments, radial fins may bepreferred. The GHR 200 is preferentially arranged vertically (i.e., witha vertically aligned longitudinal axis) with the hot ends at the top tofacilitate the filling and emptying of catalyst in the assembly outertubes 220. Addition and/or removal of catalyst may be carried out byremoval of a removable top section 209 of the vessel 205. This topsection 209 may be lifted off for servicing of the GHR 200 andreplacement of catalyst and replaced for operation of the GHR 200.

In the embodiment of FIG. 3A, the tube assemblies further include athird surrounding tube 235 that surrounds the outer catalyst tube 220.The surrounding tube 235 is sealed to the wall of the vessel 235 wall bya third tube sheet 236, and the surrounding tube 235 can be arrangedrelative to the third tube sheet 236 for engagement, such as by welding.The heating fluid 256, which can comprise, for example carbon dioxide,will flow through a space 237 defined between the outer assembly tube220 and the third surrounding tube 235 and enter an outlet cooling gascollection space 238 defined by the third tube sheet 236 and the uppertube sheet 229 before leaving as return heat fluid stream 258. The thirdsurrounding tube 235 is open ended at the hot upper end and forms a flowpath for the circulating heating fluid transferring heat to the catalystfilled space 225 between assembly outer tube 220 and the assembly innertube 220. The third surrounding tube 235 can have an inner diameter ofabout 60 mm to about 180 mm, about 70 mm to about 160 mm, or about 80 mmto about 140 mm. In an example embodiment, the third surrounding tube235 can have an inner diameter of about 102 mm.

A more detailed illustration of the arrangement of the tube assembliesis provided in FIG. 4 . The GHR 200 is illustrated relative to theexample embodiment that includes the third surrounding tube 235, but itis understood that the description otherwise also applies to the GHR 100that does not include the third surrounding tube. With reference to FIG.4 , the reaction feed stream 251, which preferably is preheated andincludes hydrocarbon and steam, enters through an inlet 250 positionedbetween the upper tube sheet 229 and the lower tube sheet 239 and passesinto the space 225 between the assembly inner tube 230 and the assemblyouter tube 222, which is filled with catalyst. The inlet flow of feedgas passes through a porous plug 233 that is positioned in the space 225near the base of the assembly outer tube 220, and the porous plug 233rests on a locating collar 234. The top portion of the assembly outertube 220 is capped so that syngas 251 a, which forms as the reactionfeed stream passes upwardly through the catalyst-filled space 225, andwhich can include some amount of unreacted hydrocarbon and steam, mustthen flow downwards through the assembly inner tube 230. The top of theassembly inner tube 230 is sealed by a porous plug 223 resting on alocating collar 224. The upper tube sheet 229 holds the assembly outertubes 220, while the assembly inner tubes 230 are held by the lower tubesheet 239. The tube sheets 229 and 239 define the feed gas inletdistribution space 208. The lower end 206 of the GHR 200 defines anoutlet collection space for the syngas product stream 253.

Located on the top of the sealed end of the assembly outer tube 220 is ashort extension tube 220 a, which is used for filling and emptyingcatalyst particles. The catalyst is filled up to a point within theextension tube 220 a well above the closed end of assembly outer tube220. The extension tube 220 a is then closed by a plug 227, which can bebiased, such as with a spring or similar mechanism, to exert a downwardforce on the catalyst particles during operation of the GHR 200. Thisarrangement prevents bed fluidization and bed movement when in operationwith reacting gas passing upwards through the catalyst bed. The topsection 207 of the pressure vessel 205 defining GHR arrangement 200 canbe removed using the flanged closure 264 to expose the tops of all ofthe assembly outer tubes 220 and their extension tubes 220 to facilitatecatalyst filling and replacement. The tube sheets 236, 229, and 239 maybe welded into the shell of the pressure vessel 205 forming the GHRarrangement 200 (or GHR 100, to which the present disclosure also fullyapplies).

As seen from the foregoing, there are a number of ways in which theshell side flow can be passed over the tube assemblies, and particularlyover the outer surface of the assembly outer tubes 120, 220. One methodis to use a baffled shell multi-pass cross-flow design, as describedabove in relation to GHR arrangement 100. Another option is to sheatheach of the tubes in an outer tube so that each tube has a purecounter-current heat transfer relationship between the reacting gases inthe catalyst bed and the heating means provided by both the coolingsyngas product delivered through the internal space of the assemblyinner tube 130, 230 and the cooling stream of carbon dioxide and waterthat is used as a heating gas around the assembly outer tube 120, 220.The outer surface of the assembly outer tubes 130, 230 can have radialfins in the case of the cross-flow arrangement and longitudinal fins forthe sheathed tube arrangement. The final selection will depend oncapital costs and pressure loss, which will increase the powerrequirement of the carbon dioxide circulation compressor. An importantconsideration in the design is to allow each individual tube to befilled with catalyst and for this catalyst to be easily removed andreplaced at regular intervals. The double tube sheet arrangement makescatalyst charge and discharge at the cooler end very difficult. Apreferred arrangement is for the reactor to be mounted vertically, withthe flow upwards through the catalyst section in the space between theassembly outer tubes 120, 220 and the assembly inner tubes 130, 230 andto fill each section from the top through a fill tube extension at theclosed end of the catalyst filled tubes. They can all be easily exposedby lifting off the top part of the reactor containment vessel. Thisarrangement with the high temperature at the top of the vessel ispreferred because it improves heat transfer from the heating fluid tothe tube assemblies while also providing the tube assemblies in anarrangement that simplifies servicing thereof, particularly in relationto removing and replacing spent catalyst material.

The oxy-fuel GHR of the present disclosure is particularly beneficial inthat it can use a high pressure heating fluid stream that is fullyoxidized and that can also have as much as 5%, 4%, 3%, 2%, or 1% oxygenon a molar basis, such as about 0.1% to about 5%, about 0.5% to about4%, or about 1% to about 2% oxygen content on a molar basis. Theprevious applications of GHR reactors coupled with either a POX or ATRsyngas generation unit were operated with a reducing fluid therein,which resulted in the shell side GHR heating gas having a highconcentration of CO, which produces carbon formation below about 850° C.in the Bouduard reaction (i.e., CO+CO=CO₂+C) and potentially severemetal dusting of the reformer tubes. The tubes must be protected byusing corrosion resistant high nickel alloy with some aluminum andchromium, such as Specialty Metals alloy 693. In addition, the tubes canbe coated with a plasma sprayed layer of alumina. All these costlyfactors can be avoided with the presently disclosed GHR configurations,particularly when combined with an oxy-fuel heating system, which ismore fully described below. A further advantage of the presentlydisclosed arrangements is that the pressure of the heating fluid stream(e.g., comprising at least carbon dioxide and water) can be at apressure that is within about 10 bar, within about 7 bar, within about 5bar, or within about 3 bar of the inlet pressure of the reactant feedstream comprising hydrocarbon and steam, which enters the catalystfilled tube assemblies. In other words, the pressure differentialbetween the heating fluid stream and the reactant stream comprising thehydrocarbon and water/steam is less than 15 bar, less than 10 bar, orless than 5 bar. The ability to operate under such a small pressuredifferential across the tube walls minimizes stress in the walls of theassembly outer tubes 120, 220, in particular, and this enables theability to utilize tubes with significantly smaller thicknesses relativeto other reforming reactors. The reduction in tube wall thicknessaccording to the present disclosure can be such that the tube walls arereduced in thickness relative to known reforming reactors by at least10%, at least 15%, at least 20%, at least 25%, at least 30%, at least40%, or at least 50%, such as a reduction in thickness of about 10% toabout 75%, about 15% to about 60%, or about 20% to about 50%.

The GHR arrangements described herein can be specifically configured foruse in the catalytic reforming of hydrocarbons, and particularlymethane, by catalytic reaction between the hydrocarbon and steam in thepresence of the catalyst. The heat of reaction and the heat required topreheat the reaction products can be provided in a variety of manners,and it is understood that the GHR arrangements are not necessarilylimited to use in combination with the further embodiments of thedisclosure. Nevertheless, the disclosed GHR arrangements areparticularly beneficial for use in the further described methods andsystems/units/plants for producing hydrogen. In some embodiments,therefore, the heat of reaction can be provided by indirect heattransfer from the circulation of a stream that comprises predominantlycarbon dioxide and that also comprises steam. The carbon dioxide plussteam stream (i.e., heating fluid stream 156, 256) can be heated by thecombustion of a hydrocarbon in oxygen. The operation of the heatingfluid stream circulating through the GHR 100, 200 at a pressure near tothe pressure of the reforming reaction itself ensures that at all pointsin the GHR arrangements 100, 200, the pressure difference between theheating fluid stream 156, 256 and the reaction feed stream 151, 251,which comprises hydrocarbon plus steam, is minimized. This enablesconfiguration of the GHR arrangements 100, 200 for operation as a hightemperature reactor system with minimal pressure differentials (as justdiscussed above), and this in turn provides the ability to use tubes andtube sheets of relatively small thicknesses (as also just describedabove), which results in an economical reactor design.

The steam plus hydrocarbon catalytic reforming reactions that produce aproduct gas mixture comprising carbon monoxide, hydrogen, and carbondioxide, together with unconverted hydrocarbon and steam, are designedto provide a reaction product temperature in the range of about 600° C.to about 1000° C., about 700° C. to about 1000° C., or about 850° C. toabout 950° C., and the reaction occurs in a space that contains agranular catalyst and that must be externally heated to achieve thenoted reaction product temperature and also since there is a very largeendothermic heat of reaction. To maximize process efficiency, it isdesirable to use the sensible heat in the reaction product gas leavingthe catalyst bed and to cool this gas down to a temperature approach ofabout 25° C. to about 100° C. relative to the preheated steam plushydrocarbon feed to the reforming reactor (GHR 100, 200). The mostconvenient way that this can be accomplished is to use a concentric tubearrangement for the GHR 100, 200 in which the catalyst is in the spacebetween the assembly outer tubes 120, 220 and the assembly inner tubes130, 230, and the reaction product stream 153, 253 comprising hydrogenplus carbon monoxide passes back along the central bore of the assemblyinner tube 130, 230. As described above, the assembly outer tubes 120,220 are capped at the hot end. The assembly outer tubes 120, 220 andassembly inner tubes 130, 230 are located on two separate tube sheets,which separately define an inlet distribution space for the preheatedreaction feed stream 151, 251 and an outlet collection space for thesyngas reaction product stream 153, 253. Both of these tube sheets areat the cooler end of the vessel 105, 205 defining the GHR 100, 200.There is no requirement for a hot end tube sheet with the concentrictube design. The individual tubes are free to expand and contract as thereactor heats up and cools down with no restraint and no effect on anyother components. The two tube sheets have minimal pressure differencesacross them, and this minimizes the required tube sheet thickness asnoted above relative to known systems. The tube sheets also are able tobe sealed into the pressure vessel 105, 205 by welding, and thiseliminates the need for expensive, very large diameter flanges to holdthe tube sheets within the vessel, as is required in known systems.

Multiple numbers of concentric tubes are used in the reactor dependingon the heat transfer medium, which will enter the shell side of thereactor containment vessel at a temperature that is about 10° C. toabout 200° C., about 20° C. to about 150° C., or about 25° C. to 100° C.higher than the temperature of the syngas reaction product 153, 253leaving the GHR 100, 200. The pressure difference between the heatingfluid stream 156, 256 and the syngas reaction product stream 153, 253leaving the GHR 100, 200 can be, in certain embodiments, in the range ofabout 1 bar to about 5 bar, about 1 bar to about 4 bar, or about 1 barto about 3 bar.

The hydrogen production methods of the present disclosure beneficiallycan incorporate the GHR arrangements 100, 200 in combination withadditional operational configurations that provide for high purityhydrogen production in a highly efficient and cost-effective manner. Insome embodiments, an oxy-fuel combustor can be used to produce theheating fluid stream 156, 256, and the fuel gas used in the oxy-fuelburner can be one or both of a waste gas from the hydrogen PSA systemand a hydrocarbon, such as methane. The heating fluid stream 156, 256can be at a pressure that is preferably less than 5 bar difference fromthe syngas reaction product stream 153, 253 leaving the GHR 100, 200 andhas a temperature that is about 25° C. to about 100° C. greater than theformed syngas entering the central bore of the assembly inner tube 130,230. The conditions for the reforming reaction that takes place in theGHR 100, 200 are such that a steam to active carbon ratio of about 3 toabout 9, about 4 to about 7, or about 5 to about 7 is used. As anexample, operation of the GHR at a pressure of about 90 bar with steamand methane as the reactants and a steam to active carbon ratio of about6 results in about 30% unreacted methane remaining in the syngasproduct. When the syngas is processed as otherwise described herein, awaste gas stream from the hydrogen PSA system can contain substantiallyall the excess methane and all the carbon dioxide formed in thesteam-methane reforming and the shift reactions. This waste gas can alsobe at about 90 bar, and can be mixed with additional methane beforebeing passed to the oxy-fuel combustor. A circulating carbon dioxidestream at about 90 bar can be mixed with the combustion products in theoxy-fuel combustor. The flowrate can be adjusted to produce the requiredtemperature for the heating fluid stream 156, 256 that enters the shellside of the GHR 100, 200.

An economizer heat exchanger can be used to recover the heat availablefrom the return heating fluid stream 158, 258 that leaves the GHRreactor at a temperature in the range of about 500° C. to about 800° C.,about 550° C. to about 750° C., or about 600° C. to about 700° C. Thisheat pre-heats the reactant feed stream 151, 251 and any additionalmethane and waste fuel gas for the oxy-fuel combustor plus preheatingthe circulating carbon dioxide stream. At the cold end of the economizerheat exchanger, the circulating oxy-fuel gas can be cooled in a directwater contact cooler where water vapor condenses, and liquid water isseparated. The product carbon dioxide stream can be removed from thecirculating carbon dioxide under pressure control, and the carbondioxide stream can enter a gas circulation compressor where its pressureis raised to a range of about 1 bar to about 10 bar, about 2 bar toabout 8 bar, or about 2 bar to about 5 bar to overcome pressure drop inthe circulating oxy-fuel system. The syngas can be produced in the GHR100, 200 at pressures in the range of about 10 bar to about 150 bar,about 15 bar to about 120 bar, or about 20 bar to about 100 bar. Thetemperature of the syngas leaving the catalyst bed can be about 600° C.to about 1000° C., about 700° C. to about 1000° C., or about 850° C. toabout 950° C. Operational parameters as described herein beneficiallycan result in a maximum pressure difference across the walls of theassembly inner tubes 130, 230 and the assembly outer tubes 120, 220 ofabout 5 bar at an operating temperature of up to 1000° C., which allowsthe tubes to be constructed from relatively thin walled stainless steel,as previously noted above.

The produced syngas enters a steam generator to produce saturated steamrequired for the reformer feed, which is then superheated in theeconomizer heat exchanger. Optionally, extra steam can be generated frompreheated boiler feed water and superheated in the economizer heatexchanger. The syngas then enters a catalytic carbon monoxide shiftconverter, which produces more hydrogen according to the reaction ofEquation 2 above. Optionally, the temperature of the syngas followingthe shift converter can be increased to a temperature in the range ofabout 450° C. to about 600° C. in the economizer heat exchanger followedby being reduced in pressure to a range of about 15 bar to about 35 bar,or about 20 bar to about 30 bar in a turbine producing power for theprocess. The heat available from the turbine exhaust gas can be used topreheat the boiler feed-water required for steam generation. The syngasis then cooled, and condensed water is separated in a direct contactwater cooler before the syngas enters the hydrogen PSA multi-bedadsorption system for hydrogen isolation. The hydrogen PSA in particularcan be paired with a membrane unit that treats the compressed hydrogenPSA waste gas to separate a hydrogen rich permeate stream withapproximately the same hydrogen concentration as the cooled syngas feedfrom the GHR 100, 200. The hydrogen PSA waste gas must be increased inpressure in a compressor to be used as a fuel gas in the oxy-fuelcombustor, and the partial pressure of the residual hydrogen in thewaste gas is high enough to allow the bulk of the hydrogen content to beremoved as a permeate stream in the membrane unit. This is recycled backto the hydrogen PSA feed, and this arrangement results in a hydrogenrecovery of greater than 98% based on the mass of the hydrogen presentin the crude hydrogen stream leaving the direct contact cooler.

As seen from the forgoing, in various embodiments, the presentdisclosure can provide a reactor for carrying out a catalytic reactionbetween a hydrocarbon fuel and steam to produce a product gas containinghydrogen and carbon monoxide. The reactor can be configured foroperation in several, various example implementations. The containmentof the catalyst can be in a vertically mounted array of tubes havingclosed ends, with a centrally located tube through which the reactionproducts are withdrawn. The reaction products passing downwards throughthe centrally located tubes can be a in heat transfer relationship withthe reacting feed gas flowing upwards through the catalyst. The catalystcan be located in a space (e.g., an annulus) defined between theassembly inner tubes and the assembly outer tubes. The catalyst can beloaded and unloaded in each tube assembly through a filling tube locatedin the top, closed end of the tube assembly. The outer surface of thecatalyst tube can be heated by a heating fluid stream flowingcounter-currently to the upward flow of the reacting gases in thesections of the tube assemblies that are filled with the catalyst. Theheating fluid can be a circulating stream comprising predominatelycarbon dioxide that has been heated by mixing with the products ofcombustion of a hydrocarbon with oxygen. The reactor is mountedvertically with its hot end at the top.

In some embodiments, each tube assembly can be surrounded by an outertube with the top inlet end open to the heating fluid which flowsdownwards around each tube assembly, thus providing heat to each tubeassembly, and the heating fluid is collected below an outlet tube sheet.In some embodiments, the heating fluid can flow downwards around thetube assemblies, providing heat to the reacting feed gases, in amulti-pass cross-flow arrangement defined by a succession of baffleplates arranged horizontally at 90 degrees to a horizontal axis of thetube assemblies. In some embodiments, each of the tubes can be sealedinto separate tube sheets, which are located at the bottom of thevertically defined reactor. The tube sheets can define the distributionspace for the feed gas and heating gas and the collection space for theproduct syngas and cooled heating fluid. In some embodiments, the outersurface of the tube assemblies can be provided with radial fins toenhance heat transfer from the heating fluid, which is in multi-passcross-flow arrangement flowing in a counter-current direction to thereacting feed gases. In some embodiments, the outer surface of the tubeassemblies can be provided with longitudinal fins along their length toenhance heat transfer from the downward flowing heating gas. In someembodiments, the reactor pressure vessel can be internally insulated. Insome embodiments, all the internal surfaces within the reactor which areexposed to partial pressures of carbon monoxide at operatingtemperatures where the Bouduard reaction can take place can be protectedfrom metal dusting corrosion by the choice of a suitable metal alloy orby a protective applied coating or by a layer of internal insulation. Insome embodiments, the maximum pressure difference between the inletheating gas and the inlet gas entering the reformer catalyst filledtubes can be about 5 bar. In some embodiments, the maximum pressuredifference between the inlet heating fluid and the inlet gas enteringthe tube assemblies can be about 2 bar. In some embodiments, thetemperature of the feed gas to the tube assemblies can be the range ofabout 400° C. to about 600° C. In some embodiments, the product gasleaving the reactor can be in the pressure range of about 20 bar toabout 100 bar. In some embodiments, the product gas leaving the catalystbed and entering the central outlet tube can be in the temperature rangeof about 800° C. to about 950° C. In some embodiments, the heating gasentering the reactor can be at a temperature of about 25° C. to about100° C. higher in temperature than the product gas leaving the catalystbed and entering the central outlet tube. In some embodiments, thecooled heating fluid leaving the GHR can be at a temperature of about25° C. to about 100° C. higher than the feed gas entering the tubeassemblies. In some embodiments, the length of the tube assemblies canbe about 6 meters to about 18 meters. In some embodiments, each of thetube assemblies can have a closed end. In some embodiments, the bottomof the space between the assembly inner tubes and the assembly outertubes can be sealed by a porous plug to allow passage of the inlet gasand confine the catalyst. In some embodiments, the top of each of theassembly inner tubes can be sealed with a porous plug through which theproduct syngas flows and which confines the catalyst. In someembodiments, an extension tube can be located on the sealed top end ofeach tube assembly and can act as a fill and empty point for thecatalyst charge in each of the tube assemblies. In some embodiments,each of the catalyst fill tubes can be provided with a spring loadedplug exerting a down-ward force on the catalyst bed in each tubeassembly.

A GHR as described herein can be utilized particularly in a hydrogenproduction unit. A hydrogen production unit specifically can include allof the components necessary for utilizing inputs of only hydrocarbon,oxygen, and water and provide outputs of substantially pure hydrogen,carbon dioxide for use or sequestration, and excess water. The combinedparts thus define a system for producing hydrogen, and the system orunit can be combined with further industrial equipment and plants sothat the produced hydrogen can be used as a feed stream into the furtherplant. The hydrogen production unit, however, may also be operatedwithout combination with other equipment or plants, and the hydrogenproduction unit may thus be operated as a stand-alone hydrogenproduction plant to export hydrogen as a product. In variousembodiments, a hydrogen production unit, system, or plant can comprisethe following: a GHR as otherwise described herein; an oxy-fuelcombustor arranged to produce a combustion product stream that can beinput to the GHR as a heating fluid; one or more shift reactors arrangedto receive a syngas product that is formed in the GHR; one or morehydrogen isolation units, which is described in greater detail below andwhich can comprise only one or more pressure swing absorbers (PSA) orwhich preferably can also comprise one or more separation membranes andone or more compressors; one or more turbines arranged to expand thesyngas product that is formed in the GHR and thus produce power with anassociated generator; and a plurality of heat exchanger members, whichcan include, for example a waste heat boiler for forming steam fromwater, a recuperator heat exchanger, water coolers, and the like. Insome embodiments, one or more pumps and/or compressors may be includedfor maintaining necessary operating pressures and flow rates in thesystem. It is likewise understood that the hydrogen production unit,system, or plant can include necessary piping, valves, and controlcomponents necessary for operation thereof.

With the above general discussion of a hydrogen production unit, system,or plant in mind, a simplified arrangement of the components thereofpresent in example embodiments of the present disclosure is provided inFIG. 5 . This illustrates a general layout of a hydrogen productionunit, system, or plant including the components discussed above. A GHR100, 200 suitable for use in such unit, system, or plant has alreadybeen described above. Additional discussion of the operation of anoxy-fuel combustor for provision of the heating fluid stream 156, 256that is formed with the oxy-fuel combustor for use in the GHR isprovided below. Also described below are the operational and likewisebeneficial aspects of utilizing a hydrogen isolation unit that includesa separation membrane and one or more compressors in combination with aPSA. Additionally, the following provides discussion of the improvedfunctionality of the hydrogen product by implementation of compressionat appropriate stages, particularly in relation to handling of the wastegas leaving the PSA, to enable combination of oxy-fuel combustion withoperation of the GHR.

In various embodiments, the present disclosure can provide a process forproduction of a syngas (e.g., hydrogen plus carbon monoxide) gas mixturein a catalytic reforming reactor using a steam plus hydrocarbon feed inwhich the endothermic heat of reaction plus the sensible heat in thereaction products is provided by heat transfer from a heating fluid. Theheating fluid particularly comprise predominately carbon dioxide.Predominately indicates that greater than 50% molar of the heating fluidis carbon dioxide. A stream comprising predominately carbon dioxide thuscan comprise at least 55%, at least 60%, at least 65%, at least 70%, atleast 75%, or at least 80% molar carbon dioxide. In some embodiments,the heating fluid stream can comprise about 55% to about 95%, about 60%to about 90%, or about 70% to about 90% molar carbon dioxide. In someembodiments, the heating fluid stream can comprise as much as 5% molaroxygen, such as about 0.1% to about 5%, about 0.5% to about 4%, or about1% to about 2% oxygen content on a molar basis. The heating fluid streammay have no more than 5%, no more than 4%, no more than 3%, or no morethan 2% oxygen molar in some embodiments. A predominately carbon dioxideheating fluid stream can comprise water, which may be liquid or may bein the form of steam. Water can comprise no more than 5%, no more than10%, no more than 15%, or no more than 20% of the heating fluid streamon a molar basis, such as about 1% to about 30%, about 2% to about 25%,or about 5% to about 20% water on a molar basis. Further, the streamcomprising predominately carbon dioxide may have a lesser water contentwhen entering the oxy-fuel combustor than when entering the reformingreactor as the heating fluid.

The stream comprising predominately carbon dioxide particularly can beheated by the combustion of a fuel, such as a hydrocarbon, with anoxidant, which may comprise substantially pure oxygen, in an oxy-fuelcombustor with direct mixing of the combustion products with the streamcomprising predominately carbon dioxide. In some embodiments, theoperating pressure in the catalytic reforming reactor can be in therange of about 25 bar to about 90 bar. In some embodiments, the molarratio of steam to the carbon in the hydrocarbon in the feed can be inthe range of about 3 to about 7. In some embodiments, the temperature ofthe steam plus hydrocarbon feeds to the catalytic reforming reactor canbe in the temperature range of about 400° C. to about 600° C. In someembodiments, the reaction products leaving the catalytic reformingreactor can be in the temperature range of about 850° C. to about 950°C. In some embodiments, the oxygen entering the oxy-fuel combustor canbe diluted with part of the circulating carbon dioxide heat transferfluid to give an oxygen concentration in the range of about 20% to about30% molar. In some embodiments, the oxygen concentration in the oxygenfeed gas can be in the range of about 90% to about 99.8% molar. In someembodiments, the oxygen required for combustion of the hydrocarbon canbe provided by an air stream from which the oxygen is separated bydiffusion through an oxygen ion transport membrane (ITM) into themixture of hydrocarbon and the circulating carbon dioxide. In someembodiments, the ITM combustor can comprise two ITM reactors operatingin series separated by two heat exchangers with the circulating carbondioxide feed first being preheated in the first heat exchanger to atemperature above about 800° C. against the circulating carbon dioxideleaving the first ITM combustor, then being heated to a temperatureabove about 800° C. in the second heat exchanger against the circulatingcarbon dioxide leaving the second ITM combustor, then being heated to atemperature of about 900° C. to 1100° C. in the second ITM combustor. Insome embodiments, the air pressure in the ITM combustor can be in therange of about 1.1 bar to about 1.5 bar. In some embodiments, the airentering the ITM combustors can be preheated against the depleted airleaving the ITM combustors in a third heat exchanger. In someembodiments, the circulating heating fluid mixture leaving the oxy-fuelcombustor and entering the catalytic reforming reactor can be at atemperature that is about 25° C. to about 100° C. higher than the syngasproduct leaving the catalyst and entering the syngas outlet tubes. Insome embodiments, the cooled circulating heating gas mixture can leavethe catalytic reforming reactor at a temperature that is about 25° C. toabout 100° C. above the temperature of the steam plus hydrocarbon feedentering the GHR tube assemblies. In some embodiments, the circulatingheat transfer fluid mixture can be cooled to near ambient temperature,and condensed liquid water derived from hydrogen in the hydrocarbon fuelto the oxy-fuel combustor can be separated. In some embodiments, thecarbon dioxide derived from combustion of the carbon in the hydrocarbonfuel in the oxy-fuel combustor can be removed from the circulating heattransfer fluid mixture at near ambient temperature following waterremoval. In some embodiments, the pressure of the circulating heattransfer fluid mixture entering the catalytic reforming reactor can beless than about 5 bar different from the pressure of the syngas productleaving the catalytic reforming reactor. In some embodiments, thecirculating heat transfer fluid mixture leaving the catalytic steam plushydrocarbon reforming reactor can be used to heat the steam plushydrocarbon feed streams to the catalytic reforming reactor inlettemperature. In some embodiments, the hydrogen plus carbon monoxideproduct stream from the catalytic reforming reaction can be cooled bytransferring sensible heat to produce steam in a water boiler, theheating being added to the hydrocarbon stream that reacts in thecatalytic reforming reactor. In some embodiments, the heating carbondioxide plus steam stream entering the catalytic reforming reactor canbe about 25° C. to about 100° C. above the temperature of the syngasstream leaving the catalyst bed and entering the assembly outlet tubes130, 230, of the GHR 100,200. In some embodiments, the syngas productfrom the catalytic reforming reactor, following cooling to form steam ina waste heat boiler, followed by carbon monoxide reacting with steam ina carbon monoxide catalytic shift reactor can be heated in an economizerheat exchanger and reduced in pressure in a power producing turbine.Optionally, a second catalytic carbon monoxide shift reactor can beadded to produce additional hydrogen in the turbine exhaust stream toform a crude hydrogen product stream followed by cooling to near ambienttemperature with condensed liquid water separation. The crude hydrogenstream can be separated into a pure, pressurized hydrogen product streamand a waste gas stream containing all the carbon dioxide derived fromthe hydrocarbon feed to the catalytic reformer and which can becompressed to a pressure required for feed to the oxy-fuel combustor. Insome embodiments, an additional stream of hydrocarbon can be added tothe total waste gas stream to provide sufficient heat release in theoxy-fuel combustor to satisfy the total heat requirement of the GHRreactor. In some embodiments, carbon dioxide can be separated by knownprocesses from the impure hydrogen stream following carbon monoxideshift conversion of carbon monoxide plus steam to hydrogen and carbondioxide. In some embodiments, the waste gas from the multi-bed PSAhydrogen separation unit can be compressed, and carbon dioxide can beremoved by known carbon dioxide separation processes. In someembodiments, the compressed waste gas, following carbon dioxide removal,can be separated in a second multi-bed PSA into a second hydrogenstream, which can be added to the first PSA hydrogen product stream tobecome the total hydrogen product stream. In some embodiments, the netproduct carbon dioxide stream separated from the circulating carbondioxide heat transfer mixture by pressure control can comprise the totalcarbon dioxide from the carbon in the hydrocarbon streams used as feedto the catalytic reformer and the oxy-fuel heater. In some embodiments,the hydrogen product can have a total impurity level below about 100 ppmmolar. In some embodiments, the hydrogen product can be used as asubstitute for natural gas in applications involving domestic,commercial, and industrial heating, as fuel for gas turbine combinedcycle power generation systems mixed with nitrogen from the cryogenicair separation unit, and as ammonia synthesis gas in a 3 to 1 ratio ofhydrogen to nitrogen. In some embodiments, the hydrogen plus carbonmonoxide gas mixture produced in the catalytic reformer can be reducedin pressure in a power producing turbine before entering the hydrogenPSA unit.

In various embodiments, the presently disclosed hydrogen production ismade particularly efficient and economical though implementation of ahydrogen isolation unit. Such unit can be implemented in a hydrogenproduction plant and process to specifically improve the ability toremove a maximum content of produced hydrogen, which is beneficial whensubstantially pure hydrogen as a product is a goal of the production(i.e., as opposed to producing impure hydrogen for use as a supplementalfuel in the process). The primary hydrogen PSA produces a substantiallypure hydrogen product stream, part of which is recycled back to the PSAfeed in order to elevate the feed stream hydrogen concentration to above70% hydrogen molar. This allows the relatively high concentration ofcarbon dioxide plus methane in the feed stream to be lowered to thepoint where a hydrogen recovery of 88% to 90% of the hydrogen in the PSAfeed can be achieved. The PSA waste gas stream, which contains about 11%of the total hydrogen feed to the PSA together with a small quantity ofcarbon monoxide and a large quantity of methane, is compressed to theoperating pressure of the SMR reactor, generally in the range of about50 bar to about 100 bar. The pressure energy present in the compressedPSA waste gas elevates the partial pressure of the hydrogen fraction.This provides an opportunity to use a gas diffusion membrane to separatea large fraction, generally in the range of about 80% to about 95%, ofthe hydrogen present in the compressed membrane feed as a diffusing lowpressure hydrogen rich permeate stream. The recovery of hydrogen isenhanced by re-cycling a small flow of pure carbon dioxide withsubstantially zero hydrogen content into the permeate side of themembrane unit counter-current to the waste gas flow. The hydrogen richpermeate stream is then compressed and recycled back to the hydrogen PSAfeed gas stream. This combination of a single hydrogen PSA, a compressedwaste gas stream and a membrane diffusion unit with the diffusing lowpressure hydrogen rich stream compressed and recycled back to thehydrogen PSA feed will result in a hydrogen recovery increased from therange 88% to 90% up to 98% to 99.5%. The hydrogen depleted compressedwaste gas stream containing 100% of the carbon derived from the SMRhydrocarbon feed is then mixed with a hydrocarbon and used as fuel inthe oxy-fuel combustor. All the carbon dioxide derived from the totalcarbon in the SMR and oxy-fuel heater hydrocarbon feed streams appearsas carbon dioxide product at the near ambient temperature part of theoxy-fuel recycle system, following liquid water separation, where it isremoved under pressure control. The result is substantially 100% carbondioxide capture with the carbon dioxide available at the same pressureas the reformer

An example embodiment of a hydrogen isolation unit is illustrated inFIG. 6 . The unit specifically provides for separation and purificationof hydrogen from a mixed gas stream, such as a syngas stream. Thehydrogen isolation unit includes a multi-bed pressure swing adsorber(PSA) unit 406, a membrane separation unit 446, a first compressor 442,and a second compressor 449. The hydrogen isolation unit provides highefficiency hydrogen production from a mixed gas feed stream thoughcombination of the PSA 406 and the membrane separation component 446since hydrogen still present in the waste gas stream leaving the PSA 406will be removed in a high concentration by the membrane separationcomponent as a permeate stream that is recycled back to the PSA 406 foroutput through the hydrogen product stream. Preferably, the membraneseparation unit 446 is effective to withdraw at least 75%, at least 80%,at least 85%, or at least 90% by weight of the hydrogen that is presentin the PSA waste gas into the permeate stream.

The multi-bed PSA unit 406 can utilize a combination of multiple feedstreams to provide the desired operational efficiency of the hydrogenisolation unit. The multicomponent feed gas stream 401 (e.g., syngas) iscombined with a compressed permeate stream 453 that contains a highproportion of the hydrogen from the PSA waste gas stream 441, as notedabove. It is also can be combined with an enhancing stream 454, which isa portion of the substantially pure hydrogen product stream 428separated in the hydrogen PSA unit 406 before the remaining portion ofthe hydrogen leaves as hydrogen product stream 456. The amount ofhydrogen in the enhancing stream 454 is chosen so that the final mixedfeed stream 451 entering the PSA 406 has a hydrogen concentration ofgreater than 60% molar, greater than 65% molar, or greater than 70%molar, such as about 60% to about 95% molar, about 65% to about 90%molar, or about 70% to about 85% molar. This enhancement enables the PSA406 to recover greater than 90%, greater than 95%, greater than 97%,greater than 98%, or greater than 99% molar of the hydrogen content inthe feed gas stream 401, in certain embodiments. The total hydrogenproduct is produced as substantially pure hydrogen stream 428. Asignificant amount of the other gases present in the multicomponent feedgas stream 401 (e.g., carbon dioxide and hydrocarbon, such as methane,plus hydrogen and carbon monoxide) will then leave in PSA waste gasstream 441.

The PSA waste gas stream 441 can be compressed in a first compressor442, preferably to a pressure that is suitable for the waste gas to beutilized for further purposes, such as delivered to a combustor in anoxy-fuel combustion process as otherwise described herein. In variousembodiments, the first compressor can be configured to compress the PSAwaste gas stream 441 to a pressure of at least 40 bar, at least 70 bar,or at least 100 bar, such as about 40 bar to about 140 bar, about 70 barto about 130 bar, or about 100 bar to about 120 bar. In certainembodiments, a pressure of about 92 bar may be preferred. The firstcompressor 442 can be a multi-stage unit with each stage followed by agas cooler, such as the intercooler 443. The compressed waste gas stream444 exiting the first compressor 442 can be cooled to near ambienttemperature in a cooler 451 producing compressed and cooled waste gasstream 445 at a temperature of about 15° C. to about 45° C., about 18°C. to about 40° C., or about 20° C. to about 30° C., such as about 25°C. in certain embodiments.

The stream 445 is fed into a membrane separation unit 446. The membraneseparation unit can comprise any equipment configured for producing apermeate stream comprising hydrogen and a non-permeate comprising theremaining portions of the entry stream. In FIG. 6 , a carbon dioxidestream 455 enters the permeate side of the membrane unit and flowscounter-current to the compressed and cooled waste gas stream 445. Thecarbon dioxide stream 455 preferably comprises less than 1000 ppmhydrogen, less than 750 ppm hydrogen, less than 500 ppm hydrogen, lessthan 250 ppm hydrogen, or less than 100 ppm hydrogen (on a molar basis).The carbon dioxide stream 455 also preferably is substantially pure,including less than 1%, less than 0.5%, less than 1%, or less than 0.1%molar of any components other than carbon dioxide. The carbon dioxidestream 455 flowing counter-current to the waste gas stream 445 enablesthe hydrogen concentration in the retentate stream 402 to be reduced toless than 5% molar, less than 4% molar, less than 3% molar, less than2%, or less than 1% molar. The diffusing permeate stream leaving theseparation membrane 446 can be at a pressure in the range of about 1 barto about 25 bar, about 1.5 bar to about 20 bar, or about 2 bar to about15 bar, and this provides for an hydrogen partial pressure differencebetween the retentate and permeate sides of the membrane 446 to cause atleast 75%, at least 80%, at least 85%, or at least 90% by weight of thehydrogen present in the cooled and compressed waste gas stream 445 todiffuse into the permeate stream 447.

The permeate product stream 447 is compressed in a second compressor 449(which can be referenced as a recycle compressor) to a pressure that issufficiently higher than the pressure of the mixed gas feed stream 401to allow the discharge flow stream 450 from the second compressor 449 topass through the cooler 542 to form stream 453 and mix with the mixedgas feed stream 401. In an example embodiments, a sufficiently highpressure may be in the range of about 10 bar to about 50 bar, about 15bar to about 45 bar, or about 20 bar to about 30 bar, such as about 26bar. The second compressor 449 can have an associated intercooler 448.The compressed permeate product stream 450 is cooled to a temperature ofabout 15° C. to about 45° C., about 18° C. to about 40° C., or about 20°C. to about 30° C., such as about 25° C. in certain embodiments in theafter-cooler 452 to form cooled and compressed permeate product stream453, which is handled as already discussed above.

As discussed in greater detail in other sections of the presentdisclosure, the retentate stream 402 leaving the membrane unit 446 canbe used for providing part of the fuel for an oxy-fuel heater. Such useis particularly beneficial when oxy-fuel heating is combined with use ofa GHR for hydrogen production. The above-described use of a membranehydrogen concentrator with recycle to the hydrogen PSA feed isparticularly enabled when the waste gas from the PSA is compressed tonear the GHR operating pressure. In other words, operation of a hydrogenisolation unit, as described here in relation to FIG. 6 , can beparticularly configured so that the first compressor 442 (see, alsocompressor 11 in FIG. 7 , as discussed below) will provide an outputstream at a pressure that is substantially close to the operatingpressure of the GHR in the associated hydrogen production system andprocess. The two pressure ranges preferably differ by no more than 20%,no more than 15%, no more than 10%, no mor than 5%, or no more than 2%.This particularly allows the residual retentate stream 402 (see, also,stream 46 in FIG. 7 , discussed below) to be at a pressure that isrequired for combustion at the GHR operating pressure. The processoperates with substantially complete carbon dioxide capture (e.g., atleast 95%, at least 98%, at least 99%, or at least 99.9% capture byweight) from the total hydrocarbon feed in the associated hydrogenproduction system and process without the requirement any additionalcarbon dioxide removal system (e.g., scrubbers, membranes, or other,known carbon dioxide capture materials). This is achieved by cooling thecirculating carbon dioxide plus water heating gas stream leaving the GHRin an economizer heat exchanger followed by a direct contact watercooler to near ambient temperature, separating condensed water, andremoving the net carbon dioxide product from the circulating carbondioxide stream under pressure control. The process also has a hydrogenrecovery from the crude hydrogen feed of greater than 98% and alsoachieves a thermal efficiency (heating value of hydrogen/heating valueof hydrocarbon feed on an LHV basis) of about 65% to about 90%, about70% to about 85%, or about 74% to about 80%. The thermal efficiencydepends on the hydrogen product pressure.

As can be seen from the forgoing, in various embodiments, a gaseous feedmixture of steam plus a hydrocarbon can be processed to produce a syngasstream comprising predominately hydrogen and carbon monoxide, togetherwith varying quantities of carbon dioxide, steam, and methane. Thesyngas can be processed to achieve conversion of the carbon monoxide tohydrogen plus carbon dioxide by catalytic reaction with steam, followedby cooling and separation of condensed water in order to produce animpure hydrogen product stream, which is separated into a substantiallypure hydrogen stream and a waste fuel gas stream. In some embodiments,the crude hydrogen can be passed into a hydrogen separation andpurification unit producing a pressurized, substantially pure hydrogenproduct stream and a low pressure waste gas stream. In some embodiments,the hydrogen separation and purification unit can comprise a hydrogenmulti-bed pressure swing adsorber producing a pressurized, substantiallypure hydrogen stream and a low pressure waste gas stream. In someembodiments, part of the substantially pure hydrogen product from thePSA can be recycled to the PSA feed giving a total feed hydrogenconcentration of about 60% to about 85% molar or about 70% to about 75%molar. In some embodiments, the waste gas can be compressed to apressure within about 15 bar, about 10 bar, about 5 bar, or about 3 barof the syngas supply pressure. In some embodiments, the compressed, nearambient temperature waste gas can be separated in a membrane gasseparator into a pressurized retentate waste gas stream and a lowpressure enriched hydrogen stream containing about 70% to about 95% ofthe hydrogen present in the hydrogen PSA waste gas stream. In someembodiments, a substantially pure carbon dioxide stream containing lessthan about 100 ppm of hydrogen can be passed through the permeate sideof the membrane counter-current to the waste gas. In some embodiments,the waste gas stream from the permeate side of the membrane can be at apressure about 2 bar to about 10 bar. In some embodiments, the waste gasstream from the permeate side of the membrane can be compressed andrecycled to the hydrogen PSA feed. In some embodiments, the hydrogenplus carbon monoxide gas mixture can be supplied at pressure in therange of about 30 bar to about 100 bar. In some embodiments, thehydrogen recovery in the system can be greater than about 98% of thehydrogen in the crude hydrogen that is fed to the PSA plus membranesystem. In some embodiments, the process can comprise producing the feedmixture comprising a catalytic steam plus hydrocarbon tubular reformerconvectively heated by a carbon dioxide recycle gas which has beenheated by direct mixing with the combustion products of the waste fuelgas from the retentate side of the membrane unit plus added hydrocarbonburning in substantially pure oxygen. In some embodiments the syngasfollowing high temperature carbon monoxide shift reaction can beexpanded to a lower pressure in a power producing turbine. In someembodiments, the substantially pure hydrogen from the PSA can be mixedwith nitrogen preheated in the hydrogen plant and used as fuel gas for agas turbine. In some embodiments, the substantially pure hydrogen fromthe PSA can be mixed with nitrogen as feed to an ammonia synthesis unitwith heat from the hydrogen plant plus heat from the ammonia synthesisreaction providing the heat for a power unit. In some embodiments, thepower system burning hydrogen can use a closed cycle carbon dioxideworking fluid with a power producing turbine, a recuperator heatexchanger, and a recycle carbon dioxide compressor. In some embodiments,the use in combination of hydrogen production and hydrogen separationcan be accomplished with approximately 100% carbon dioxide capture basedon carbon in the total hydrocarbon feed to the process.

An example embodiment of a hydrogen production system and the associatedmethod of operation of the system is illustrated in FIG. 7 , the systemincluding a GHR, oxy-fuel combustion, and hydrogen isolation accordingto the one or more embodiments of the disclosure. For clarity, the GHRis an example embodiment of a reforming reactor that is useful in themethods and systems for oxy-fuel hydrogen combustion. Thus, reference tothe GHR herein is understood to implicitly include reference to areforming reactor more broadly, unless the context expressly dictatesotherwise.

With specific reference to FIG. 7 , a catalytic tubular GHR arrangement1 (having internal configurations consistent with the disclosure alreadyprovided herein in relation to GHR 100 and 200) has an inlet flowthrough line 60 into the feed gas inlet distribution space (element 208,FIG. 3A) between the upper tube sheet (element 229, FIG. 3A) and thelower tube sheet (element 239, FIG. 3A) so that the feed gas enters thecatalyst filled space in the tube assemblies of the GHR 1. The inletflow provides a reaction feed stream comprising steam and hydrocarbon,which preferably is a low carbon number gas, such as C₁ to C₅hydrocarbons (individually or in a mixture), and particularly methane ornatural gas. The steam and hydrocarbon preferably are provided in amolar ratio of about 2 to about 9, and preferably about 4 to about 7,such as a ratio of about 5 in certain embodiments. The reactant feedstream in line 60 preferably is at a pressure of about 30 bar to about120 bar, and preferably about 30 bar to about 100 bar, such as apressure of about 90 bar in certain embodiments. The temperature of thereactant feed stream can be in the range of about 300° C. to about 700°C., and preferably about 400° C. to about 600° C., such as a temperatureof about 550° C. in certain embodiments.

The syngas (predominately carbon monoxide and hydrogen) that forms bythe catalytic reaction of the hydrocarbon and steam in the sections ofthe GHR 1 tube assemblies filled with catalyst can initially have atemperature of about 600° C. to about 1000° C., preferably about 800° C.to about 1000° C., and more preferably about 850° C. to about 950° C.,such as a temperature of about 900° C. in certain embodiments, whenentering the assembly inner tubes (elements 130, 230, FIG. 2A and FIG.3A) and then passes down the central bore of the assembly inner tubesand leaves the GHR arrangement 1 as stream 25 at a temperature of about400° C. to about 800° C., preferably about 500° C. to about 700° C.,such as a temperature of about 600° C., in certain embodiments. Inparticular, the syngas can be cooled in the assembly inner tube to atemperature that is within about 25° C. to about 100° C. of thetemperature of the reactant feed stream. The syngas formed in GHR 1 canhave a carbon monoxide concentration that is dependent upon the steam tocombined carbon ratio of the hydrocarbon used and the temperature of thesyngas leaving the catalyst and will preferably be in the range of about2% to about 15%, about 2.5% to about 10%, or about 3% to about 8% on amolar basis, such as a concentration of about 5% molar in certainembodiments.

Stream 25 is cooled in a steam generator 4, which is further equippedwith a steam separator 3 and a thermo-syphon boiler feed-watercirculation unit with circulating streams 27 and 26. The cooled syngasleaves the steam generator 4 in stream 57 at a pressure of about 15 barto about 120 bar, preferably about 20 bar to about 115 bar, and morepreferably about 28 bar to about 98 bar, such as a pressure of about 88bar in certain embodiments, and a temperature of about 200° C. to about400° C., preferably about 225° C. to about 350° C., such as about 313°C. in certain embodiments. The cooled syngas then enters a catalyticcarbon monoxide shift reactor 5 where its carbon monoxide content isreduced. Preferably, the carbon monoxide shift reactor is effective toreduce carbon monoxide content in the syngas stream to no greater than2%, no greater than 1%, or no greater than 0.5% molar. In this exampleembodiment, the carbon monoxide content is reduced from 5.09% to 0.27%molar. The temperature of the shifted syngas leaving in stream 56 willincrease in temperature by about 10° C. to about 80° C. or about 25° C.to about 70° C., such as to a temperature of about 365° C. in thisexample embodiment. The outlet syngas stream 56 from the shift reactor 5and the saturated steam stream 65 exiting the steam separator 3 areseparately passed through the economizer heat exchanger 2 at anintermediate point and are both heated to a temperature of about 400° C.to about 650° C., about 450° C. to about 625° C., or about 500° C. toabout 600° C., such as a temperature of about 550° C., in certainembodiments. The heated syngas leaving the economizer heat exchanger instream 58 is expanded in a power producing turbine 7 driving anelectrical generator 8 to a pressure of about 10 bar to about 40 bar,about 15 bar to about 35 bar, or about 20 bar to about 32 bar, such as apressure of about 27 bar in certain embodiments. The turbine outletstream 55 of the shifted and heated syngas stream is cooled in heatexchanger 23 to a temperature of about 25° C. to about 55° C., about 30°C. to about 50° C., or about 35° C. to about 45° C., such as about 40°C. in certain embodiments. This heats a boiler feed-water stream 28received from pump 24 to a temperature of about 225° C. to about 375°C., about 250° C. to about 350° C., or about 275° C. to about 325° C.,such as about 300° C. in certain embodiments, before it enters the steamseparator 3 as stream 28. Alternatively, the turbine 9 driving theelectrical generator 10 can be eliminated so that the high pressuresyngas stream 41 can be cooled in heat exchanger 2, and the hydrogen athigh pressure can be retained in the hydrogen product stream leaving thePSA 21. The boiler feed-water stream 28 from pump 24 can be heated inthe economizer heat exchanger 2. The cooled syngas then enters a directwater cooled heat exchange packed column 15 to be further cooled so thatcondensed water leaves from the bottom of the column. The condensedwater splits into water stream 31, which joins the boiler feed-watersource stream 29 for input into pump 24 and water stream 32, whichcirculates through heat exchanger 16 by pump 20 to leaves as waterstream 33 for input back into column 15. Heat is removed from the systemwith a cooling water stream 34 to 35 in heat exchanger 16. The totalwater stream entering pump 24, which includes the make-up water flow 29,is purified in a purifier such as an ion exchange unit, which is notshown.

The cooled syngas exits the packed column 15 as stream 43, saturatedwith water vapor at a pressure of about 20 bar to about 35 bar, about 20bar to about 32 bar, or about 24 bar to about 30 bar, such as a pressureof 26 bar in certain embodiments, and at a temperature of about 15° C.to about 40° C., about 20° C. to about 35° C., or about 22° C. to about28° C., such as about 25° C. in certain embodiments. Stream 43 joinswith an impure hydrogen recycle stream 42 and a pure hydrogen recyclestream 70 to form combined gas stream 44, which is passed into amulti-bed hydrogen PSA (pressure-swing absorption) unit 21. The PSA unit21 provides a substantially pure (e.g., less than 1000 ppm, less than500 ppm, or less than 100 ppm impurity concentration) hydrogen productstream at a pressure of about 15 bar to about 35 bar, about 20 bar toabout 30 bar, or about 22 bar to about 28 bar, such as about 25 bar incertain embodiments and a temperature of about 15° C. to about 35° C.,about 20° C. to about 30° C., and about 22° C. to about 28° C., such asabout 25° C. in certain embodiments. Part of the pure hydrogen productcan be taken as hydrogen side stream 70 for being recycled back to thePSA feed stream 44 to raise the hydrogen concentration to preferablygreater than 70% molar to facilitate the recovery of the large quantityof carbon dioxide and methane present in the waste flow from the PSAwhile maintaining high purity hydrogen recovery from the PSA. A wastegas exits the PSA unit 21 as waste gas stream 67, which is compressed incompressor 11 driven by a motor 12 to a pressure of about 30 bar toabout 100 bar, about 40 bar to about 100 bar, or about 80 bar to about100 bar, such as a pressure of about 95 bar in certain embodiments. ThePSA waste gas exits the compressor 11 as stream 68 and is processedthrough a membrane separation unit 22, such as a polymeric membraneunit. The membrane separation unit 22 can be any membranous uniteffective to separate the waste gas stream into a permeate stream, whichincludes residual hydrogen and may include additional impurities, and aretentate stream, which comprises the remaining components of the PSAwaste gas stream, such as carbon dioxide, water, and carbon monoxide. Asmall, substantially pure carbon dioxide stream 71 (e.g., comprisingless than 100 ppm hydrogen content) is fed into the permeate side of themembrane 22 at the retentate exit end of the membrane to promotediffusion of hydrogen to achieve the lowest partial pressure of hydrogenin the final waste product flow stream 46. The permeate stream 45 (i.e.,impure hydrogen) is provided at a pressure of about 5 bar to about 15bar, about 7 bar to about 14 bar, or about 8 bar to about 12 bar, suchas a pressure of about 10 bar in certain embodiments. This stream isthen compressed by a factor of about 1.5 to about 4 or about 2 to about3 to achieve a pressure of about 18 bar to about 50 bar, about 20 bar toabout 40 bar, or about 22 bar to about 32 bar, such as about 26 bar incertain embodiments in compressor 13, which is driven by a motor 14. Thecompressed permeate stream is provided as recycle hydrogen stream 42,which is added to the cooled syngas stream 43 exiting the column 15. Therecycle hydrogen stream 42 can have approximately the same hydrogen molfraction as the cooled syngas stream 43. This recycle system can beeffective to provide a hydrogen recovery of greater than 98% of thehydrogen contained in cooled syngas stream 43. As already noted above,the combination of PSA and membrane separation can define a hydrogenisolation unit or system together with the compressors 11, 13 so that amaximum proportion of the hydrogen in the shifted syngas in cooledsyngas stream 43 can be isolated as the hydrogen product stream. Ahydrogen isolation unit such as this can be effective to recover atleast 95%, at least 97%, or at least 98% molar of the hydrogen in thecooled syngas stream 43, such as about 95% to about 99.9%, about 96% toabout 99.8%, or about 97% to about 99.5% molar of the hydrogen. Thehydrogen isolation unit is further described herein in relation to FIG.6 .

A hydrocarbon fuel source 47 provides a hydrocarbon fuel stream 49 foraddition to the retentate portion stream 46 of the waste gas from thePSA 21 and separation membrane 22. The hydrocarbon fuel stream 47 can beat a pressure of about 70 bar to about 120 bar, about 80 bar to about110 bar, or about 90 bar to about 100 bar, such as a pressure of about94.6 bar in certain embodiments to form the total fuel gas stream 50,which is heated in the economizer heat exchanger 2 to a temperature ofabout 400° C. to about 700° C., about 450° C. to about 650° C., or about500° C. to about 600° C., such as about 550° C. in certain embodimentsbefore entering the oxy-fuel combustor 6 as stream 62. There it isburned with oxygen from substantially pure oxygen stream 63, which hasbeen preheated to a temperature of about 200° C. to about 300° C., about215° C. to about 285° C., or about 225° C. to about 275° C., such asabout 250° C. in certain embodiments using a portion of the saturatedsteam stream 65. In alternative embodiments, a diluted oxygen stream maybe used. For example, oxygen stream 63 may be blended with a portion ofthe recycle carbon dioxide stream 52 to form an oxidant stream having anoxygen concentration of about 15% to about 75%, about 17% to about 50%,or about 20% to about 30% on a molar basis, the remainder substantiallycomprising only carbon dioxide. The diluted oxidant stream then can beheated in the economizer heat exchanger 2 to a temperature of about 450°C. to about 700° C., about 475° C. to about 650° C., or about 500° C. toabout 600° C., such as about 530° C. in certain embodiments.

The combustor 6 can have any configuration recognized as useful in anoxy-fuel combustion process. For example, the combustor can be arrangedso as to define an outer combustor shell 6 a and a combustor liner 6 bthat defines internally a combustion chamber 6 c. The fuel and oxidantmay be injected into the combustor 6 into the combustion chamber 6 c.Oxidant may also be injected through at least a portion of the liner 6b. The stream comprising predominately carbon dioxide likewise can beinjected through the combustor liner. The oxy-fuel combustor 6 can bearranged to receive a first part of the stream comprising predominatelycarbon dioxide into a reaction zone 6 e of the combustion chamber 6 cand to receive a second part of the stream comprising predominatelycarbon dioxide into a dilution zone 6 d of the combustion chamber 6 c.It is understood in FIG. 7 that the combustor is not illustrated toscale and rather illustrates the presence of the component parts, whichcan be arranged as needed. A suitable combustor further is described inU.S. Pat. No. 10,859,264, the disclose of which is incorporated hereinby reference. Additional combustor arrangements are described in U.S.Pat. No. 9,068,743, the disclosure of which is also incorporated hereinby reference.

The combustion products mix with a heated recycle carbon dioxide stream61 that has been heated in the economizer heat exchanger 2 to atemperature of about 400° C. to about 700° C., about 450° C. to about650° C., or about 500° C. to about 600° C., such as about 550° C. incertain embodiments providing the outlet heating fluid stream 64 at atemperature of about 750° C. to about 1150° C., about 800° C. to about1100° C., or about 1000° C. to about 900° C., such as about 950° C. incertain embodiments, which enters the top, or high temperature end, onthe shell side of the GHR arrangement 1. The heating fluid stream 64 canhave a carbon dioxide concentration of about 85% to about 98% molar,about 90% to about 98% molar, or about 92% to about 96% molar. Theremaining fraction is preferably water, but small amounts of otherimpurities may be present. In particular, the heating stream 64 can havean oxygen concentration of about 0.1% to about 5%, about 0.2% to about3%, or about 3% to about 2% on a molar basis.

The heating fluid stream 64 passes downwards in the GHR around the outersurfaces of the tube assemblies, particularly around the outer surfacesof the assembly outer tubes (120, 220) as facilitated by the baffles(155 in FIG. 2A) or the space between the assembly outer tubes and thethird surrounding tubes (235 in FIG. 3A). The fluid cools as heat isprovided to the tube assemblies, and it leaves the GHR 1 as returnheating fluid stream through line 59 at a temperature of about 500° C.to about 800° C., about 550° C. to about 750° C., or about 600° C. toabout 700° C., such as a temperature of 658° C. in the present exampleembodiment. It then enters the economizer heat exchanger 2 and exits asstream 53 after being cooled to a temperature of about 20° C. to about80° C., about 25° C. to about 70° C., or about 30° C. to about 60° C.,such as about 40° C. in certain embodiments. The stream 53 is passed toa purification unit. In the illustrated embodiment, the purificationunit comprises a direct contact packed tower water cooler 17; however,it is understood that the purification unit can comprise furthercomponents that are effective for separating carbon dioxide from theother components of the heating fluid from the reforming reactor aftercooling in the economizer heat exchanger. The water cooler 17 has acirculation system where water in stream 36 is circulated to stream 39with a pump 19 and a water cooled heat exchanger 18 with cooling waterflow 37 to 38. Excess water leaves in stream 30 to combine with boilerfeed-water source stream 29. The cooled water-saturated carbon dioxidestream has a saturation water mol fraction in the range of about 0.002to about 0.004 at 25° C. according to experimental results. Leaving thetop of the water cooler is a cooled and purified stream comprisingpredominately carbon dioxide, stream 40. This stream 40 is compressed ina circulation compressor 9 driven by a motor 10 to a pressure of about75 bar to about 110 bar, about 80 bar to about 100 bar, or about 88 barto about 92 to leave as stream 41, which splits into product and recyclestreams. In this manner, the stream of predominately carbon dioxide 41is at a pressure suitable for passage through the oxy-fuel combustor 6.The net carbon dioxide product stream 51 is removed under pressurecontrol, and the recycle carbon dioxide stream 52 enters the economizerheat exchanger 2 to be heated and passed back to the combustor 6 asdiscussed above. The hydrocarbon feed stream from hydrocarbon source 47is preferably at a pressure of about 75 bar to about 115 bar, about 80bar to about 110 bar, or about 85 bar to about 105 bar, such as about 95bar in certain embodiments, is sent in part as reformer feed stream 48,which will be sent to the GHR 1, but a portion (stream 49) is dividedout for addition to the PSA waste in stream 46 to form fuel stream 50,which will be combusted in the combustor 6. Both of streams 48 and 50are heated to a temperature of about 400° C. to about 700° C., about450° C. to about 650° C., or about 500° C. to about 600° C., such asabout 550° C. in certain embodiments in the economizer heat exchanger 2.The reformer feed stream 48 leaves the economizer heat exchanger 2 asheated reformer feed stream 67 and is mixed with the superheated steamstream resulting from the steam stream 65 having been heated in theeconomizer heat exchanger, as already noted above. The mixed,steam-saturated fuel stream 60 preferably has a steam to fuel ratio ofabout 7 to 1 to about 2 to 1, about 6 to 1 to about 3 to 1, or about 5to 1 in certain embodiments. The mixed, steam-saturated fuel stream 60then enters the tube assemblies in the GHR arrangement 1 for catalyticreaction in the catalyst-filled spaces between the assembly outer tubes(120, 220) and assembly inner tubes (130, 230).

Calculations for operation of the above-described system are based on100 lb mols methane feed to the GHR arrangement. Total methane feed at95 bar is 125.42 lb mols. Hydrogen production at 25 bar is 313.7 lbmols. Oxygen required at 95 bar is 94.93 lb mols. There is excess heatavailable of 2200 Kw from a temperature of 167° C. down to 40° C., whichcan be used for district heating or boiler feedwater preheating in anassociated steam power system, particularly when the hydrogen is used topower a combined cycle gas turbine power unit. carbon dioxide productionat 94 bar is 125.42 lb mols.

While the foregoing provides discussion of preferred embodiments of thedisclosure, particularly including all of the GHR assembly, oxy-fuelcombustion, hydrogen isolation and compression, and other featuresbeneficial for providing high efficiency hydrogen production, thepresent disclosure also encompasses other aspects whereby hydrogenproduction can also be achieved at desirable efficiencies and may bebetter utilized in certain circumstances. An example embodiment,therefore, of an additional hydrogen production process according to thepresent disclosure is provided below with reference to FIG. 8 .

In FIG. 8 , an oxy-fuel hydrogen production system is illustrated.Although the example embodiment that follows is described utilizingspecific operating parameters, it understood that such operatingparameters relate only to a preferred embodiments, and such parametersare subject to variation within ranges, particularly ranges otherwisedescribed herein for operation in similar embodiments of oxy-fuelhydrogen production.

The system and process for oxy-fuel hydrogen production that isillustrated in FIG. 8 can preferably produce hydrogen at a pressure ofabout 28 bar. Note that the GHR 50 is shown inverted for simplicity. Theoxy-fuel hydrogen production process includes a low temperature carbondioxide removal step on the pressurized syngas after cooling to nearambient temperature in the direct water contact tower 56. This limitsthe carbon dioxide concentration in the PSA feed and allows a highenough hydrogen content in the pressurized vent gas stream 91 leavingthe carbon dioxide removal unit 68 for the stream 91 to be processed ina second small PSA unit 69 producing an additional pure hydrogen productstream 89 and a final waste gas stream 88. The separate removal ofcarbon dioxide and hydrogen from the waste stream 88 means that it canbe re-compressed and recycled back to the GHR catalytic reformer withadditional hydrocarbon feed.

With reference to FIG. 8 , the GHR 850 has a feed stream 897 at about 32bar and about 550° C. comprising a mixture of fresh methane stream 895plus waste gas stream 888 from a second stage PSA 869 which has beencompressed to about 32 bar in compressor 870 with motor drive 871. Thetwo streams are separately heated in heat exchanger 854, the waste gasto a temperature of about 140° C., which is about 5° C. colder than thedew-point of the cooling syngas product stream 884, which passes fromthe low temperature shift reactor 855 into heat exchanger 854 at about257° C. The methane stream is heated to a temperature of about 250° C.The two exit streams 8110 and 896 are heated to about 550° C. in heatexchanger 877 against the cooling circulating carbon dioxide heattransfer stream 8116, which enters heat exchanger 877 at about 620 167C. A stream of methane 8100 is separated from the 550° C. methane streamand enters the oxy fuel combustor 878 together with the 25% oxygen plus75% carbon dioxide molar oxidant stream 898 and the carbon dioxidecirculating stream 899 which have both been heated in 877 to about 509°C. The heated circulating carbon dioxide stream 8115, which is a mixtureof combustion products and the circulating carbon dioxide stream, entersthe GHR 850 at about 950 167 C. The remaining methane plus the heatedwaste gas streams then mix with the 32 bar, 550° C. steam stream 8109,which has been produced from the waste heat boiler (WHB) 851 and thesteam separator 852 as stream 8108 and superheated in heat exchanger877.

The product syngas stream 879 at about 600° C. leaving the GHR is cooledto about 330° C. in the WHB 851 leaving as stream 880 to enter the hightemperature shift reactor 853 where its temperature rises to about 425°C. and then cooling to about 257° C. in a second pass through the WHB851 leaving as stream 882 to enter heat exchanger 854. The syngas coolsto about 210° C. then leaves as stream 883 and enters the lowtemperature shift reactor 855 before leaving at a temperature of about260° C. as stream 884, re-entering heat exchanger 854. The syngas leavesheat exchanger 854 as stream 885 at about 40° C. and is cooled to about25° ° C. in the direct contact water cooler 856, which has a circulatingwater pump 858 and a cooling water heat exchanger 857. The outlet waterstream 8120 derived from the oxidation of hydrogen in the total methanefeed stream 895 is sent to a water treatment system 860. The outletstream 886 is separated in the first PSA unit 861 into a hydrogenproduct stream 887 at about 28 bar and a 1.2 bar waste gas stream 8118which is raised to about 40 bar pressure in compressor 863 driven bymotor 862. The compressed waste gas stream 8119 is dried in a dual bedthermally regenerated drier 867 using nitrogen gas, then stream 8120enters the low temperature carbon dioxide removal unit 868. Theseparated carbon dioxide streams 892 enter the multi-stage compressor864 with intercooling together with the carbon dioxide from the oxy-fuelcombustion of methane, stream 8121. The product stream from thecompressor at about 75 bar is cooled to near ambient temperature bywater cooler 894, and the resulting high density supercritical carbondioxide is increased in pressure to about 200 bar in a multi-wheelcentrifugal pump 866, delivering the carbon dioxide product stream 893for pipeline transportation. The uncondensed waste gas stream 891 atabout 39 bar pressure containing about 67.6% hydrogen is separated inthe second PSA unit 869 into a second hydrogen product stream 889 atabout 38 bar and a waste gas stream 888 at about 1.2 bar, which entersthe compressor 870.

The cooled circulating carbon dioxide steam 8117 leaving heat exchanger877 at about 40° C. is cooled to 25° C. in the direct contact watercooler 874, which has circulating water pump 875 and a cooling waterheat exchanger 876. The cooled carbon dioxide stream 8102 is compressedin the gas circulation compressor 872 driven by motor 873. The netproduct carbon dioxide from the oxy-fuel combustion is removed as stream8121 from the compressor discharge stream 8103. This stream divides intoa portion which, with added oxygen stream 8111, form the oxidant stream8112 entering the heat exchanger 877. Part of this, stream 8113, isheated to about 140° C., stream 8114, in heat exchanger 854 againstcooling syngas product and then rejoins the circulating oxidant streamin heat exchanger 877 where the combined stream is heated to about 550°C. The net condensed water from the direct contact coolers 856 and 874,streams 8120 and 8162 plus the water feed stream 8103 are purified inthe water treatment unit 860. The total pure water stream 8105 is pumpedto about 35 bar pressure in the boiler feed-water pump 859, heated inthe heat exchanger 854 to about 232° C. as stream 8107, and sent to thesteam separator 852 as boiler feed water for the WHB 851.

The syn-gas stream 884 from the low temperature shift converter 855cools in heat exchanger 854 to its dew-point at a temperature of about150° C., which is a pinch point temperature in the heat exchanger. Thesyngas cooling from about 150 C results in a very large additionalamount of heat being available as the 15.9% of residual steam content inthe syngas condenses. Additionally, there is a similar amount of heatavailable from the cooling carbon dioxide circulating heat transferstream, with about 6% steam content condensing below its dew pointtemperature of about 117° C. in heat exchanger 877. Taking these twostreams together, this heat can be used for heating a circulatingpressurized water flow to about 133° C., which is an ideal temperaturelevel for district heating to replace natural gas or oil currently used.An alternative or additional use would be to preheat the hydrogenproduct to about 133° C. prior to its use as fuel in either new orexisting gas turbines converted to burn a 60% hydrogen plus 40% nitrogenfuel gas or for ammonia production.

In some embodiments, air can be utilized as the oxidant for combustingthe hydrocarbon fuel and waste gas provided into the combustor.Likewise, a gas turbine exhaust stream, which generally contains about11% to about 13% oxygen and is available at a temperature of about 400°C. to about 650° C., may be used as the oxidant. This can eliminate theneed for an ASU and substitutes a gas turbine, which acts as a toppingcycle on the combustor and actually increases the process efficiency forthe hydrogen production. A gas turbine can be integrated with acatalytic steam plus hydrocarbon reformer to produce a hydrogen product.The reformer can be a GHR arrangement as already described above. In theGHR arrangement, the syngas product formed by catalytic reaction ofsteam and hydrocarbon fuel flows down the assembly inner tubes (elements130 and 230 in FIG. 2A and FIG. 3A, respectively), which extend to theend of the catalyst filled outer tubes, in counter-current heat transferrelationship with the catalyst filled tubes to provide part of therequired heat for the endothermic reactions. The remaining heat must beprovided by a heating fluid which passes over the outer catalyst filledtube.

In each case considered, the gas turbine fuel provided is a hydrocarbon,such as methane or LPG or light naptha, at a pressure required by thechosen gas turbine. The substantially pure, pressurized hydrogen productproduced by the reformer process is delivered from a multi-bed pressureswing adsorption unit, which also produces a low pressure waste fuelgas, which contains all the carbon in the reformer hydrocarbon feed,predominantly as carbon dioxide, but also with some unconverted carbonmonoxide and methane. This waste fuel is typically at a pressure ofabout 1.1 bar to about 1.3 bar. This low pressure fuel gas can be burnedusing oxygen contained in the gas turbine exhaust to produce atemperature in the burner exhaust in the range of about 1000° C. toabout 1200° C. The quantity of waste fuel gas burned reduces the oxygencontent of the burner exhaust to about 3% to about 6% of the oxygencontent in the exhaust leaving the gas turbine. The combustion of thePSA waste gas, which will contain all the carbon derived from themethane reformer feed, together with additional methane to use all theavailable oxygen present in the gas turbine exhaust, will result in ahigh temperature gas comprising carbon dioxide plus nitrogen and steamwith about 1% to about 2% oxygen. The burner exhaust has a carbondioxide concentration of about 20% to about 30% compared to the 3% to 4%content in the gas turbine exhaust. The waste fuel gas can be burned inthe gas turbine exhaust in a grid type burner to uniformly heat the gasturbine exhaust. This arrangement can utilize the waste fuel at thedelivery pressure of the hydrogen PSA unit, or very close to thispressure. The quantity of waste fuel gas burned defines the maximumhydrogen production of the reformer unit since the maximum quantity ofavailable oxygen in the gas turbine exhaust has been used to support thecombustion.

In the first case considered, the heat delivered by the heated gasturbine exhaust is used to heat a recycle carbon dioxide stream which,in turn, provides heat for the reformer. In the second case, the heatedgas turbine exhaust is used directly as the heating fluid in thecatalytic steam plus hydrocarbon reformer. In each case the gas turbineexhaust, following the reformer heating, is used to provide heat bothfor preheating the hydrocarbon feed gas for the reformer feed, the wastefuel gas, producing the excess steam required for the reformer, and tosuperheat the total steam flow to the reformer. The gas turbine exhaust,following reformer heating, also produces steam for power production andfor heating in the amine carbon dioxide removal unit, which receives thegas turbine exhaust leaving the heat exchange section. The heatingmedium is the heat produced by burning the PSA waste gas plus optionallyadditional hydrocarbon fuel to consume the remaining oxygen in the hotgas turbine exhaust leaving only about 1% to about 2% residual oxygenconcentration. In the first case, the hot combustion gas is used to heata pressurized circulating carbon dioxide stream to a temperature in therange of about 900° C. to about 1050° C. The circulating carbon dioxidestream provides the heat for the catalytic steam plus methaneendothermic reactions in the gas heated reformer plus preheat for thereactants. The pressure of the circulating carbon dioxide stream can bein the range of about 25 bar to about 100 bar. The pressure is withinabout 5 bar of the pressure at the inlet to the catalyst filled reformertubes.

An example embodiment illustrating the first case is shown in FIG. 9 .The illustration provides the details of the gas turbine, the use of thewaste gas as fuel in a burner, and the circulating fluid stream that isheated to provide the heating fluid stream for the reforming reactor.Further elements useful in such embodiments may be immediatelyrecognized in light of the additional disclosure herein, such as inrelation to the systems and methods described in relation to FIG. 7 andFIG. 8 , as well as the further discussion of the GHR arrangementsabove. Likewise, operating parameters described below reference theillustrated example embodiment, but it is understood that parametersalready described herein in terms of ranges would likewise apply to theexample embodiment discussed below. For example, input and outputpressures and temperatures for the reforming reactor may be in any ofthe ranges already described above.

With reference to FIG. 9 , a gas turbine is shown receiving an inlet airflow 371 with a turbine section 356 driving a compressor section 354 andan electric generator 372 and with a combustor 355 receiving apressurized hydrocarbon fuel stream 334. The gas turbine chosen for thisexample is a Siemens SGT-800 which has a net power at iso conditions of62.5 Mw with an LHV efficiency of 41.4% an exhaust temperature of about596° C. and an exhaust flow of about 135.5 Kg/second, but it isunderstood that such equipment is described for illustrations purposesof this example embodiment. The fuel used in this example is methane,but other fuels may likewise be used. The gas turbine exhaust 342 entersa combustor 307 (e.g., a grid burner), which extends across the gasturbine exhaust duct in which a fuel gas stream 337 at a pressure ofabout 1.1 bar is burned using oxygen in the gas turbine exhaust, whichenters at a concentration of about 11.6% molar in this example, althoughother concentrations as already noted above may be utilized. About 95%of the oxygen is consumed in the burner 307.

The gas turbine is integrated with a reforming reactor 301. The fuel gasstream 337 is the waste gas from a multi-bed pressure swing hydrogenseparation unit 336, which processes the cooled crude hydrogen stream341 from the reformer 301. Such hydrogen separator 336 may be arrangedas already described herein. A substantially pure hydrogen productstream 390 and water stream 391 exit the separator along with a wastegas stream 392. The waste gas from the hydrogen separator contains allthe carbon present in the hydrocarbon feed to the reformer,predominantly as CO₂ but also present in unconverted hydrocarbon andcarbon monoxide, and this carbon appears in the burner exhaust 343 ascarbon dioxide together with the carbon dioxide produced in the gasturbine combustor 355. The burner exhaust 343 enters a heat recoveryheat exchanger 308 at a temperature of about 1100° C. and heats a closedcycle heating gas carbon dioxide stream 339, which enters at about 270°C. and about 32 bar and leaves as stream 340 at about 950° C. Stream 340enters the reforming reactor 301 as a heating fluid that functions asalready described above to provide the heat required to perform theendothermic reforming reactions and leaves the reactor as stream 382 atabout 625° C. and about 31 bar. Stream 382 of the heating fluid leavingthe reforming reactor 301 cools in the heat exchanger 305, providing theheat for preheating any number of additional streams as alreadydescribed herein.

The recycle carbon dioxide stream 382 cools to about 270° C. and leavesthe heat exchanger 305 as stream 338 at about 31.25 bar and enters acirculation compressor 324 driven by a motor 325. The heat exchanger 305provides the preheating for the fuel gases and steam requirements. Thehydrogen separator waste fuel gas stream 326, for example, is heated toabout 550° C. and leaves as stream 337 before entering the duct burner307. In some embodiments, extra heat can be provided by diverting aportion of the heated recycle carbon dioxide stream 340 adding it to therecycle carbon dioxide stream 382 before it enters the heat exchanger305. The gas turbine exhaust stream 344 can be process for furthertreatment as needed and as would be understood in the power productionfield.

In the first case, then, it can be seen that a heating gas can beprovided in relation to heat transfer from a deoxidized hot dischargegas from a gas turbine to a circulating carbon dioxide stream as analternative to heating the circulating carbon dioxide by combusting fuelin pure oxygen at high pressure. In particular, the heating fluid can bea circulating stream of predominately carbon dioxide that can be heatedby indirect heat transfer by the exhaust from a gas turbine, whichexhaust has been increased in temperature by using the residual oxygenin the gas turbine exhaust as an oxidant to combust a fuel gascomprising the low pressure waste gas from the PSA, optionally combinedwith additional hydrocarbon. The cooled gas turbine exhaust can betreated as necessary for carbon dioxide removal.

The second case uses the hot combustion gas to directly heat thecatalyst tubes. This means a much higher pressure difference across thecatalyst tube wall than the first case, which limits the operatingpressure at the inlet to the catalyst tubes to a range of about 25 barto about 35 bar and means that thick walled high nickel tubes such asHK40 must be used. This makes the direct use of the deoxidized gasturbine exhaust as a heating gas far less attractive than the use of therecycle high pressure carbon dioxide heating gas heated by the gasturbine exhaust. It enables the use of the GHR compact catalyticreformer already described, which will have a far lower cost than thedirectly heated reformer, which has a low pressure heating gas with amuch lower heat transfer coefficient resulting in much greater heattransfer area required and with expensive catalyst tubes. As such, thedirect heating case is less preferred. In each case, the hot gas turbineexhaust following the reformer heating is used to preheat the steam andhydrocarbon feed streams to the reformer and to provide heat requiredfor regeneration of the chemical solvent used in the carbon dioxiderecovery unit.

Although non-preferred, the direct heating of the GHR is a possibleroute for providing oxy-fuel hydrogen production, and such directheating route can be carried as follows. A gas turbine can receive aninlet air flow with a turbine section driving a compressor section andan electric generator and with a combustor receiving a pressurizedhydrocarbon fuel stream 334. The gas turbine may be, for examples aSiemens SGT-800, which has a net power at iso conditions of 62.5 Mw withan LHV efficiency of 41.4% an exhaust temperature of 596° C. and anexhaust flow of 135.5 Kg/second (e.g., using methane as a fuel). The gasturbine exhaust enters a grid burner, which extends across the gasturbine exhaust duct in which a fuel gas stream at a pressure of 1.1 baris burned using oxygen in the gas turbine exhaust, which enters at aconcentration of 11.6% molar. Approximately 95% of the oxygen isconsumed in the burner. The gas turbine is integrated with a catalyticsteam plus methane tubular gas heated reformer (GHR) as alreadydescribed herein. The fuel gas stream is the waste gas from a multi-bedpressure swing hydrogen separation unit, which processes the cooledcrude hydrogen stream from the reformer. The product reformed hydrogenplus carbon monoxide gas stream at 30 bar 650° C. leaving the reformerreactor enters a processing section comprising waste heat boilerproducing near saturated high pressure steam, two stage carbon monoxideshift reactors, heat exchangers for heat recovery, final direct contactwater cooler plus condensed water separator stream followed by the PSAunit producing the substantially pure hydrogen product stream. Suchcomponents may be selected from materials that are recognized as usefulin catalytic reforming processes producing substantially pure hydrogen.The carbon monoxide catalytic shift reactors convert about 95% of thecarbon monoxide in the syngas produced in the reformer reactor byreaction with excess steam producing hydrogen plus carbon dioxide. Thewaste gas from the PSA contains substantially all the carbon present inthe methane feed to the reformer, predominantly as carbon dioxide butalso present in unconverted methane and CO, and this carbon appears inthe burner exhaust as carbon dioxide together with the carbon dioxideproduced in the gas turbine combustor and carbon dioxide produced fromthe combustion of an additional methane fuel gas stream taken from themethane feed stream.

The burner exhaust enters a section of the heat recovery heat exchangerat a temperature of about 1100° C. The inlet temperature is reduced toabout 1100° C. by means of a recycle quench gas stream, which is takenfrom the gas turbine exhaust as stream and increased in pressure in theblower. The first section heats a closed cycle heating gas carbondioxide stream, which enters at about 270° C. and 32 bar and leaves atabout 950° C. before entering a catalytic steam plus methane reformerwhere it provides the heat required to perform the endothermic reformingreactions and preheat the reaction products leaving the reactor at about625° C. and 31 bar. The recycle stream then cools in the economizer heatexchanger providing the heat for preheating the methane feed andsuperheating the steam feed and the feed water stream to produce themixed the reformer methane plus water feed stream at about 31 bar andabout 550° C. plus heating the PSA waste gas stream and the additionalmethane combustion stream to about 550° C.

The design of the gas heated tubular reformer reactor with itsconcentric tube arrangement and three tube sheets has been describedabove and illustrated in FIG. 2A, FIG. 2B, FIG. 3A, FIG. 3B, and FIG. 4. These three tube sheets define, in order, a space for the collectionof the syn-gas product, a space for the inlet of the combined methaneplus water stream at about 31 bar and about 550° C. with a molar ratioof water to methane of about 4 to 1 and a space for the outletcirculating carbon dioxide stream at about 31 bar and about 625° C.,which enters a multi-channel compact heat exchanger. The recycle carbondioxide stream cools to about 270° C. leaving about at about 31.25 aboutbar and enters a circulation compressor driven by a motor. The heatexchanger provides the preheating for the fuel gases and steamrequirements. The heat exchangers and carbon monoxide shift reactors,converting carbon monoxide by reaction with excess steam to carbondioxide and hydrogen produce the following feed streams which arepreheated in heat exchanger. The methane total feed steam is at about 33bar and about 250° C. and leaves at about 550° C. This stream iscombined with the superheated steam stream at about 550° C. to form thefeed stream to the reformer with a molar ratio of steam to methane ofabout 4 to 1.

The PSA waste fuel gas stream plus additional methane stream provide atotal stream at about 1.5 bar and about 250° C. and is heated to about550° C. before entering the duct burner. The boiler feed water at about232° C., which produces a portion of the total superheated steam flow,which is mixed with heated methane to provide the feed stream for thereformer 301. The heat required for these preheating duties in the heatexchanger is greater than the heat available in the cooling recyclestream leaving the reformer. The extra heat is provided by diverting aportion of the heated recycle carbon dioxide stream at about 950° C. andadding it to the recycle carbon dioxide stream before it enters the heatexchanger. Alternatively, the stream can be taken as a side-stream fromthe heating carbon dioxide recycle stream in the heat exchanger at about625° C.

The gas turbine heat exchanger section has an inlet stream of gasturbine exhaust at about 270° C., which contains the combustion productsfrom the duct burner. It heats a feed water stream and generates astream 37 at a pressure of about 3.5 bar. This provides the steam flowto which is added the steam production from the hydrogen plant requiredfor the total 3.5 bar steam flow for regeneration duty in the aminecarbon dioxide separation unit. The gas turbine exhaust stream at about100° C., leaving the heat recovery heat exchanger, enters a directcontact water cooler where it is cooled to about 20° C., and the bulk ofthe water formed in the gas stream is condensed and removed fortreatment. The dried gas is compressed from about 0.95 bar to about 1.5bar in the compressor driven by a motor. At least 95% of the carbondioxide in the gas turbine exhaust, which is 95% of the carbon in thetotal feed to the system, is removed as a substantially pure carbondioxide stream at an average pressure (two streams at differentpressures are available) of about 2 bar and the final waste gas streamis discharged to the atmosphere. The carbon dioxide product iscompressed to about 70 bar in a multistage compressor driven by themotor 314 and leaves to enter a cooler where its temperature is reducedto about 25° C. The pressure is reduced to about 6 bar in a valveproducing a saturated liquid carbon dioxide product stream at about 6bar and a recycle flash gas stream at about 6 bar, which is recycled tothe compressor.

In one or more embodiments, the present disclosure can provide foroxy-fuel hydrogen production incorporating a gas turbine in combinationwith a catalytic reformer (GHR) using steam plus hydrocarbon as areactant stream. The reformer catalyst filled tubes can be heated by anexternal gas flow. At least part of the heat required to heat theexternal gas flow can be produced by burning the waste gas from amulti-bed pressure swing adsorber, which purifies the hydrogen productgas from the reformer in the gas turbine exhaust. The oxygen requiredfor combustion of the PSA waste gas can be taken from the exhaust from agas turbine. The waste gas from the PSA can be burned in the gas turbineexhaust, optionally with additional methane. The oxygen content of thegas turbine exhaust after combustion can be about 1% to about 2% molar.The heated gas turbine exhaust can be used to heat a circulating gasstream, which heats the catalyst filled reformer tubes in the GHR. Thedeficiency in the heat available when using only the waste hydrogen PSAwaste gas burning with the maximum available quantity of oxygen presentin the gas turbine exhaust compared to the heat required by the hydrogenplant can be remedied by adding additional hydrocarbon to fuel gas usedin the combustor. The carbon dioxide present in the gas turbine exhaustleaving the heat exchangers can be separated in a physical or chemicalabsorption process recovering at least 95% of the carbon dioxide presentin the gas turbine exhaust. Any heat required for solvent regenerationin the carbon dioxide removal process can be provided using heat presentin the deoxidized gas turbine exhaust.

In one or more embodiments, hydrogen production according to the presentdisclosure can be achieved with additional increased efficiency byimplementing options that can eliminate the need for high energy inputcomponents, such as an air separation unit (ASU) for generatingsubstantially pure oxygen for the oxy-fuel combustor that is used toproduce the heating fluid stream for the GHR arrangement. In someembodiments, this can be achieved by use of one or more ion transportmembrane units.

The general arrangement of an ITM oxy-fuel combustor and carbon dioxiderecycle heater is shown in FIG. 10 . Oxygen for the combustion isprovided from a preheated air stream at near atmospheric pressure bydiffusion of the oxygen molecules through an oxygen ion transportmembrane (ITM) operating in a temperature range of about 800° C. toabout 1050° C. An inlet flow of a circulating recycle carbon dioxidestream 1212 (corresponding to stream 61 in FIG. 7 ), mixes with a part1228 of a preheated hydrocarbon gas plus total waste fuel gas stream1226 (corresponding to stream 62 in FIG. 8 ). Both of these streams areat a temperature in a range of about 400° C. to about 700° C., about450° C. to about 650° C., or about 500° C. to about 600° C., such asabout 550° C. in certain embodiments and a pressure in the range ofabout 15 bar to about 1340 bar, about 20 bar to about 120 bar, or about25 to about 100 bar.

The mixed stream 1227 enters a heat exchanger 1202 where it is heated toa temperature of about 700° C. to about 1000° C., about 725° C. to about900° C., or about 750° C. to about 850° C., such as about 800° C. incertain embodiments. The outlet stream 1213 enters a first ITMcombustion unit 1207. The noted inlet temperature ensures that a rapiddiffusion of oxygen occurs through the oxygen ion transport diffusionmembrane 1210. The diffusing oxygen comes from an air stream 1221 at apressure of about 1 bar to about 2 bar, such as about 1.5 bar, and atemperature of about 600° C. to about 900° C., about 625° C. to about800° C., or about 650° C. to about 750° C., such as about 700° C. incertain embodiments, which is part of the heated air stream 1219 leavingthe recuperative air heat exchanger 1203. The air stream 1221 enters theair side of the first ITM combustion unit 1207. It is heatedconvectively by heat transfer through the ITM membrane 1210 as it passesthrough the unit 1207. Oxygen diffuses through the ITM 1210, and thedepleted air stream 1222 leaves the first ITM combustor air side at atemperature of about 800° C. to about 1000° C., such as about 875° C. incertain embodiments. The stream 1222 will have been depleted of about50% to about 80%, about 60% to about 80%, or about 70% to about 80%molar of the contained oxygen, which will have diffused through themembrane 1210. The diffusing oxygen reacts with the hydrocarbon (e.g.,methane) in the recycle carbon dioxide stream, which can be at apressure of up to about 50 bar up to about 75 bar, or up to about 100bar, and the heat produced raises the temperature of the mixture ofcarbon dioxide recycle flow and carbon dioxide plus water combustionproducts.

The heated recycle carbon dioxide stream 1214 leaves the first ITMcombustor 1207 at a temperature of about 800° C. to about 1100° C.,about 850° C. to about 1050° C., or about 900° C. to a about 1000° C.,such as about 950° C. in certain embodiments and enters the heatexchanger 1202 where it cools to a temperature of about 25° C. to about100167 C greater than the inlet temperature of stream 1213, and itleaves as stream 1236. This carbon dioxide recycle stream enters heatexchanger 1230 where it is heated to a temperature of about 650° C. toabout 950° C., about 700° C. to about 900° C., or about 750° C. to about850° C., such as about 800° C. in certain embodiments and exits asstream 1235. Heating is taken from the outlet carbon dioxide recyclestream 1216 leaving the second ITM combustor 206 at a temperature ofabout 800° C. to about 1300° C., about 900° C. to about 1200° C., orabout 950° C. to about 1150° C., such as about 1050° C. in certainembodiments, which itself is cooled to a temperature of about 800° C. toabout 1100° C., about 850° C. to about 1050° C., or about 900° C. toabout 1000° C., such as about 950° C. in certain embodiments, and itleaves heat exchanger 1230 as exit stream 1234, which is equivalent tostream 64 in FIG. 7 .

Stream 1235 mixes with the remaining portion of the hydrocarbon andwaste fuel gas feed stream 1229, and the mixed stream 1215 enters thesecond ITM combustor 1206. The diffusing oxygen for combustion comesfrom an air stream 1220 at a pressure of about 1 bar to about 2 bar,such as about 1.5 bar, and a temperature of about 600° C. to about 900°C., about 600° C. to about 800° C., or about 650° C. to about 750° C.,such as about 700° C. in certain embodiments, which is part of theheated air stream 1219. This air stream enters the air side of the ITMcombustor 1206 and the depleted air stream 1223 leaves at a temperatureof about 900° C. to about 1100° C., about 850° C. to about 1050° C., orabout 900° C. to about 1000° C., such as about 950° C. in certainembodiments. The recycle air stream heats by convective heat transferthrough the ITM membrane 1211 as it passes through the ITM membrane unit1206. The depleted air stream 1223 leaving the air side of the secondITM combustor 1206 is joined by stream 1222 from the first ITM combustor1207, and the total stream 1224 enters the air recuperative heatexchanger 1203. There it cools to a temperature of about 40° C. to about110° C., about 50° C. to about 100° C., or about 60° C. to about 90° C.,such as about 75° C. in certain embodiments and is vented to theatmosphere as stream 1225.

An adiabatic air compressor 1204 driven by a motor 205 with an inlet airflow 1207 delivers an air stream 1218 at a pressure of about 1.2 bar toabout 5 bar, about 1.2 bar to about 3 bar, or about 1.2 bar to about 2.5bar, such as about 1.7 bar in certain embodiments and at a temperatureof about 35° C. to about 95° C., about 45° C. to about 85° C., or about55° C. to about 75° C., such as about 65° C. in certain embodiments. Theair stream 1218 is sent to the air recuperative heat exchanger 1203where it is heated to a temperature of about 500° C. to about 900° C.,about 600° C. to about 800° C., or about 650° C. to about 750° C., suchas about 700° C. in certain embodiments. The recycle carbon dioxidestream leaving as stream 1234 then enters the GHR arrangement in any ofthe further embodiments of a hydrogen production system that aredescribed herein to function as the heating fluid stream. Theperformance of the two stage ITM combustion system can be maximized byoperation of a control system in which the hydrocarbon and air flows toeach ITM combustor 1206, 1207 are flow controlled to achieve thespecified temperatures in the system.

As can be seen from the forgoing, the oxy-fuel combustion aspect ofhydrogen production as described herein can be provided in a variety ofmanners. Combustion can specifically be carried out in a conventionalcombustor into which a hydrocarbon fuel is injected to be combusted withoxygen that has been separated from air to avoid introducing anynitrogen into the system. Combustion is carried out in the presence of astream of carbon dioxide to produce a combustion exhaust streamcomprising predominately carbon dioxide with a content of water andpotentially small fractions of one or more impurities. This type ofoxy-fuel combustion is described above in relation to the hydrogenproduction systems and methods illustrated in FIG. 5 , FIG. 7 , and FIG.8 . It is thus understood that such figures and their related disclosureexpressly describe oxy-fuel combustion in a conventional combustor withpurified oxygen to form a heating fluid stream that is configured to beadded to a GHR arrangement for heating the tube assemblies in whichhydrocarbon fuel and steam are catalytically reacted to form syngas.Such figures and their related disclosure are not, however, limited toonly the use of the conventional combustor that is described. Rather, itis expressly understood that the conventional combustor (e.g., thelabeled unit in FIG. 5 , element 6 in FIG. 7 , and element 78 in FIG. 8) may be replaced with an ITM unit as described above in relation toFIG. 10 . In such embodiments (e.g., a combination of the ITM unit ofFIG. 10 with any of the oxy-fuel hydrogen production systems of FIG. 5 ,FIG. 7 , or FIG. 8 ) the need for an ASU is eliminated since plain aircan be utilized as the oxygen source input to the oxygen side of the ITMcombustion units 1206, 1207 and since the oxygen will selectivelydiffuse across the ITM membranes 1210, 1211 in the ITM combustion units1206, 1207 to combust the hydrocarbon fuel passed through the fuel sidesof the units. When implementing the ITM combustion unit, powerconsumption required for operating the hydrogen production plant can bereduced by at least 25%, at least 35%, or at least 45%, such as about25% to about 75%, about 35% to about 65%, or about 45% to about 55%relative to an identical hydrogen production plant using a conventionalcombustor and an ASU to provide the purified oxygen.

A typical ITM membrane suitable for use as discussed above may comprisemixed metal oxides arranged in a perovskite crystal structure. A mixturecomprising carbon dioxide and a gaseous hydrocarbon fuel, such asmethane and PSA waste gas, can be passed through the permeate side ofthe ITM membrane at high pressure in the range of about 20 to about 100bar. Diffusion will take place due to the extremely low equilibriumpartial pressure of oxygen on the permeate side with oxygen reactingwith the hydrocarbon at the ITM operating temperature. The result is aprocess with 100% capture of carbon dioxide derived from the carbonpresent in the hydrocarbon feeds to both the syngas generation and theoxy-fuel combustion.

Hydrogen produced as described herein can be delivered for a variety ofend uses. As a non-limiting example, hydrogen mixed with nitrogen can beused as a fuel, replacing or supplementing natural gas, in a gas turbinecombined cycle power unit. The molar fuel gas composition can be, forexample, about 60% hydrogen plus about 40% nitrogen. Performance can bebased on the production of hydrogen at about 25 bar pressure and about25° C. with purity of about 99.995% mixed with nitrogen with a purity ofabout 99.999%. Conversion of a GE combined cycle power system comprisingtwo GE 9HA-02 gas turbines with a single steam system and using the sameheat input for the hydrogen fuel gas as the natural gas, for followingconditions are achieved: net power output (iso conditions) is about 1681Mw; heat rate is about 5306 Btu/Kwhr Net; Hydrogen required is about826756 Nm³/hr; nitrogen required is about 551171 Nm³/hr; oxygen requiredis about 8486.9 Mt/day; carbon dioxide production (methane fuel) isabout 15573.6 Mt/day; efficiency is about 51.14%. Note that the oxygenrequirement is 0.2104 Metric tons/Mwhr compared to a supercriticalcarbon dioxide cycle power requirement of about 0.54 Mt/Mwhr.

As a further non-limiting example of hydrogen use, the hydrogen producedaccording to the present disclosure can be blended with nitrogen to givean approximate 25% nitrogen plus 75% hydrogen synthesis gas at about 25bar and about 25° C. for ammonia production. The total impurity in thesynthesis gas can be less than about 50 ppm. The ammonia loop canoperate with no purge gas bleed from the loop eliminating the purge gaspurification system. The excess low grade heat below about 300° C.available from the hydrogen system can be used with the excess heatproduced in the ammonia synthesis reactor to provide heat for powerproduction. Power can be produced by transferring this heat to a highpressure circulating carbon dioxide stream that is then passed through apower producing turbine. The high temperature turbine exhaust can becooled in a recuperator heat exchanger then compressed to the turbineinlet pressure and heated in the recuperator heat exchanger. Theadvantage of using a carbon dioxide working fluid rather than steam isthe absence of a temperature plateau as boiler feed water is evaporatedin a steam system. The oxy-fuel hydrogen system can have a heat recoveryeconomizer heat exchanger that has a significant quantity of heatavailable below a temperature pinch in the range of about 250° C. toabout 300° C. This excess heat is available for preheating thecompressed 75% hydrogen plus 25% nitrogen syngas feed to the ammonialoop and also for preheating the recycle carbon dioxide. A significantquantity of excess power is generated after satisfying all the powerrequirements of the ammonia plant. This results in a lower total naturalgas requirement for this process than any other disclosed ammoniaproduction system. Performance can be as shown in the following exampleembodiment: feed gas is 250,000 Nm³/hr hydrogen plus 83,333 Nm³/hrnitrogen; the total impurity level can be about 50 ppm. The gas streamis at about 25 bar and is preheated in the hydrogen oxy-fuel system toabout 167° C.; ammonia production can be about 3039 Mt/day; excess powerproduction can be about 63.8 Mw; heat equivalent with power efficiencyof 60% can be about 25.52 million Btu/Mt of NH₃.

The terms “about” or “substantially” as used herein can indicate thatcertain recited values or conditions are intended to be read asencompassing the expressly recited value or condition and also valuesthat are relatively close thereto or conditions that are recognized asbeing relatively close thereto. For example, unless otherwise indicatedherein, a value of “about” a certain number or “substantially” a certainvalue can indicate the specific number or value as well as numbers orvalues that vary therefrom (+ or −) by 5% or less, 4% or less, 3% orless, 2% or less, or 1% or less, and any one of such values may be usedinterchangeably with the words “about” and/or “substantially” as neededfor clarity. Similarly, unless otherwise indicated herein, a conditionthat substantially exists can indicate the condition is met exactly asdescribed or claimed or is within typical manufacturing tolerances orwould appear to meet the required condition upon casual observation evenif not perfectly meeting the required condition. In some embodiments,the values or conditions can be defined as being express and, as such,the term “about” or “substantially” (and thus the noted variances) canbe excluded from the express value.

Many modifications and other embodiments of the presently disclosedsubject matter will come to mind to one skilled in the art to which thissubject matter pertains having the benefit of the teachings presented inthe foregoing descriptions and the associated drawings. Therefore, it isto be understood that the present disclosure is not to be limited to thespecific embodiments described herein and that modifications and otherembodiments are intended to be included within the scope of the appendedclaims. Although specific terms are employed herein, they are used in ageneric and descriptive sense only and not for purposes of limitation.

1. An oxy-fuel heated, hydrogen production system comprising: areforming reactor arranged to receive a stream comprising a hydrocarbonand water through a first inlet and separately receive a stream of aheating fluid through a second inlet, the reactor including a catalystcomponent effective for catalyzing a reaction between the hydrocarbonand the water to form a synthesis gas stream comprising at leasthydrogen and carbon monoxide, and the reactor including a synthesis gasoutlet arranged for exit of the synthesis gas stream from the reformingreactor; an oxy-fuel combustor arranged to receive a fuel, an oxidant,and a stream comprising predominately carbon dioxide and comprising acombustor outlet for exit of a combustion product stream from theoxy-fuel combustor, the oxy-fuel combustor being configured to combustat least a portion of the fuel with oxygen from the oxidant to formcarbon dioxide and water, which is combined with the stream comprisingpredominately carbon dioxide to form the combustion product stream; ahydrogen isolation unit arranged to receive at least a portion of thesynthesis gas stream, and provide at least part of the hydrogen from thesynthesis gas stream as a substantially pure hydrogen product stream;and a purification unit arranged to receive at least a portion of thecombustion product stream and output a stream of substantially purecarbon dioxide, the purification unit also being arranged to deliver atleast a portion of the substantially pure carbon dioxide as the streamcomprising predominately carbon dioxide; wherein the reforming reactorand the oxy-fuel combustor are functionally configured so that at leastpart of the combustion product stream is provided through the secondinlet of the reforming reactor as the stream of the heating fluid. 2.The oxy-fuel heated, hydrogen production system of claim 1, wherein thereforming reactor comprises a pressure containment vessel and at leastone set of concentrically arranged tubes positioned within the pressurecontainment vessel, each of the at least one set of concentricallyarranged tubes comprising: an outer catalyst tube; an inner reactionproduct gas tube; and catalyst material positioned within a spacedefined between an inside surface of the outer catalyst tube and anoutside surface of the inner reaction product gas tube.
 3. The oxy-fuelheated, hydrogen production system of claim 2, wherein the at least oneset of concentrically arranged tubes positioned within the pressurecontainment vessel are arranged vertically so that an upper end of theat least one set of concentrically arranged tubes defines a hot endwhere the reforming reactor operates with a higher temperature, and alower end of the at least one set of concentrically arranged tubesdefines a cold end where the reforming reactor operates with a lowertemperature, relative to the hot end.
 4. The oxy-fuel heated, hydrogenproduction system of claim 3, wherein the reforming reactor furthercomprises an upper tube sheet that is arranged to functionally alignwith the outer catalyst tube, and a lower tube sheet that is arranged tofunctionally align with the inner reaction product gas tube.
 5. Theoxy-fuel heated, hydrogen production system of claim 4, wherein thereforming reactor is arranged so the first inlet opens into a spacedefined between the upper tube sheet and the lower tube sheet.
 6. Theoxy-fuel heated, hydrogen production system of claim 4, wherein thereforming reactor is arranged so that the stream comprising ahydrocarbon and water entering through the first inlet passes upwardly,from the cold end toward the hot end, through the space within which thecatalyst material is positioned.
 7. The oxy-fuel heated, hydrogenproduction system of claim 4, wherein the reforming reactor is arrangedso that the synthesis gas outlet is positioned at a level of thereforming reactor that is below a position of the first inlet.
 8. Theoxy-fuel heated, hydrogen production system of claim 4, wherein thereforming reactor is arranged so that the synthesis gas outlet ispositioned below the lower tube sheet.
 9. The oxy-fuel heated, hydrogenproduction system of claim 4, wherein a bottom of the lower tube sheetand a bottom portion of the pressure containment vessel define acollection space for the synthesis gas stream, which proceeds downwardlyfrom the hot end through an inner bore of the inner reaction product gastube.
 10. The oxy-fuel heated, hydrogen production system of claim 3,wherein the second inlet of the reforming reactor is configured toreceive the heating fluid in an arrangement so that the heating fluidprovides heat to the outer catalyst tube.
 11. The oxy-fuel heated,hydrogen production system of claim 10, wherein the arrangement is suchthat the heating fluid entering the second inlet of the reformingreactor contacts the hot end of the at least one set of concentricallyarranged tubes and flows downwardly around an outer surface of the outercatalyst tube toward a second outlet through which the heating fluidleaves the reforming reactor.
 12. The oxy-fuel heated, hydrogenproduction system of claim 11, wherein the second outlet is positionedat a level of the reforming reactor that is above a position of thefirst inlet.
 13. The oxy-fuel heated, hydrogen production system ofclaim 11, wherein the reforming reactor further comprises a surroundingtube positioned around the at least one set of concentrically arrangedtubes, the surrounding tube being arranged to form heating spacerelative to the at least one set of concentrically arranged tubes anddefine a flow path of the heating fluid through the heating space. 14.The oxy-fuel heated, hydrogen production system of claim 11, wherein thereforming reactor further comprises a plurality of baffles attached toan inner surface of the pressure containment vessel and arranged todirect flow of the heating fluid for contact with the at least one setof concentrically arranged tubes.
 15. The oxy-fuel heated, hydrogenproduction system of claim 3, wherein the upper end of the at least oneset of concentrically arranged tubes defines a filling tube with aremovable plug.
 16. The oxy-fuel heated, hydrogen production system ofclaim 15, wherein the removable plug is configured to provide biasedforce toward the catalyst within the at least one set of concentricallyarranged tubes.
 17. The oxy-fuel heated, hydrogen production system ofclaim 3, wherein an outer surface of the at least one set ofconcentrically arranged tubes comprises a plurality of fins configuredto facilitate heat transfer between the heating fluid and the at leastone set of concentrically arranged tubes.
 18. The oxy-fuel heated,hydrogen production system of claim 3, wherein at least a portion ofinternal surfaces of the reforming reactor that are exposed to a partialpressure of carbon monoxide at operating temperatures where a Bouduardreaction occurs are protected from metal dusting corrosion by thepresence of a protective coating or a layer of internal insulation. 19.The oxy-fuel heated, hydrogen production system of claim 2, wherein thespace defined between the inside surface of the outer catalyst tube andthe outside surface of the inner reaction product gas tube that isfilled with catalyst defines a section having a length about 6 meters toabout 18 meters.
 20. The oxy-fuel heated, hydrogen production system ofclaim 1, further comprising at least one shift reactor configured toconvert at least a portion of the carbon monoxide in the synthesis gasfrom the reforming reactor to carbon dioxide and output a shift streamcomprising at least hydrogen, carbon dioxide, and waste gas.
 21. Theoxy-fuel heated, hydrogen production system of claim 20, wherein thehydrogen isolation unit comprises an inlet arranged to receive the shiftstream, output a pressurized stream of substantially pure hydrogen, andoutput a stream comprising at least part of the waste gas.
 22. Theoxy-fuel heated, hydrogen production system of claim 21, wherein thehydrogen isolation unit comprises a hydrogen multi-bed pressure swingadsorber (PSA) configured to output the pressurized stream ofsubstantially pure hydrogen and output the stream comprising at leastpart of the waste gas.
 23. The oxy-fuel heated, hydrogen productionsystem of claim 22, wherein the PSA is configured with a hydrogenrecycle line arranged to send part of the pressurized stream ofsubstantially pure hydrogen back to the inlet of the PSA.
 24. Theoxy-fuel heated, hydrogen production system of claim 22, wherein thehydrogen isolation unit further comprises at least one compressorarranged to receive and compress at least a portion of the streamcomprising at least part of the waste gas and output a compressed wastegas stream.
 25. The oxy-fuel heated, hydrogen production system of claim24, wherein the hydrogen isolation unit further comprises a membrane gasseparator having an inlet arranged to receive the compressed waste gasstream, and wherein the membrane gas separator is configured to separatethe compressed waste gas stream into a pressurized retentate waste gasstream and a hydrogen-enriched permeate stream.
 26. The oxy-fuel heated,hydrogen production system of claim 25, wherein the membrane gasseparator comprises an inlet arranged to receive a stream ofsubstantially pure carbon dioxide for passage through a permeate side ofa membrane in the membrane gas separator counter-current to thecompressed waste gas stream.
 27. The oxy-fuel heated, hydrogenproduction system of claim 25, wherein the hydrogen isolation unitfurther comprises a recirculation line through which thehydrogen-enriched permeate stream is passed back to the inlet of thePSA.
 28. The oxy-fuel heated, hydrogen production system of claim 25,further comprising a line through which at least part of the pressurizedretentate waste gas stream is passed to the oxy-fuel combustor.
 29. Theoxy-fuel heated, hydrogen production system of claim 22, wherein theoxy-fuel heated, hydrogen production system further comprises a gasturbine.
 30. The oxy-fuel heated, hydrogen production system of claim29, further comprising a line through which at least a portion of thepressurized stream of substantially pure hydrogen is passed to the gasturbine.
 31. The oxy-fuel heated, hydrogen production system of claim22, wherein the oxy-fuel heated, hydrogen production system furthercomprises an ammonia synthesis unit.
 32. The oxy-fuel heated, hydrogenproduction system of claim 22, further comprising a line through whichat least a portion of the pressurized stream of substantially purehydrogen is passed to the ammonia synthesis unit.
 33. The oxy-fuelheated, hydrogen production system of claim 1, further comprising apower producing turbine arranged to receive at least a portion of thesynthesis gas stream and expand said stream for power production. 34.The oxy-fuel heated, hydrogen production system of claim 1, wherein theoxy-fuel combustor comprises an outer combustor shell and a combustorliner that defines internally a combustion chamber.
 35. The oxy-fuelheated, hydrogen production system of claim 34, wherein the oxy-fuelcombustor is arranged to receive at least part of the stream comprisingpredominately carbon dioxide through the combustor liner.
 36. Theoxy-fuel heated, hydrogen production system of claim 34, wherein theoxy-fuel combustor is arranged to receive a first part of the streamcomprising predominately carbon dioxide into a reaction zone of thecombustion chamber and to receive a second part of the stream comprisingpredominately carbon dioxide into a dilution zone of the combustionchamber.
 37. The oxy-fuel heated, hydrogen production system of claim 1,wherein the oxy-fuel combustor comprises an ion transport membrane (ITM)combustor.
 38. The oxy-fuel heated, hydrogen production system of claim37, wherein the ITM combustor comprises an oxygen ion transportdiffusion membrane separating an air side of the ITM combustor from afuel side of the ITM combustor.
 39. The oxy-fuel heated, hydrogenproduction system of claim 38, wherein the oxygen ion transportdiffusion membrane is effective to draw oxygen from air passing throughthe air side of the ITM combustor into the fuel side of the ITMcombustor for combustion of fuel passed through the fuel side of the ITMcombustor.
 40. The oxy-fuel heated, hydrogen production system of claim37, wherein the oxy-fuel heated, hydrogen production system comprises aplurality of ITM combustors.
 41. The oxy-fuel heated, hydrogenproduction system of claim 1, further comprising a heat exchangerarranged to receive at least a portion of the heating fluid after theheating fluid exits the reforming reactor and configured to transferheat from the heating fluid to one or more further streams.
 42. Theoxy-fuel heated, hydrogen production system of claim 41, wherein the oneor more further streams to which the heat is transferred from theheating fluid include one or more of the fuel that is received by theoxy-fuel combustor, the oxidant that is received by the oxy-fuelcombustor, the stream comprising predominately carbon dioxide that isreceived by the oxy-fuel combustor, and the stream comprising thehydrocarbon and water that is received by the reforming reactor.
 43. Theoxy-fuel heated, hydrogen production system of claim 41, furthercomprising a purification unit arranged to receive the heat fluid afterleaving the heat exchanger and configured to output the streamcomprising predominately carbon dioxide.
 44. The oxy-fuel heated,hydrogen production system of claim 43, further comprising a compressorarranged to receive the stream comprising predominately carbon dioxideleaving the purification unit and configured to compress the streamcomprising predominately carbon dioxide to a pressure suitable for inputto the oxy-fuel combustor.